Carbo-metallic oil conversion with liquid water

ABSTRACT

A process is disclosed for the production of high octane gasoline and/or other valuable lower molecular weight products from carbo-metallic oils. Examples include crude oil, topped crude, reduced crude, residua, the extract from solvent de-asphalting and other heavy hydrocarbon fractions. These carbo-metallic oils contain quantities of coke precursors and heavy metal catalyst poisons substantially in excess of what is normally considered acceptable for FCC processing (fluid catalytic cracking) and substantial amounts of sulfur, nitrogen and other troublesome components may also be present. Such carbo-metallic oils are converted to the desired products in a catalytic conversion process. Named &#34;RCC&#34; (Reduced Crude Conversion) after a particularly common or useful carbo-metallic feed, the present process is by no means restricted to reduced crude or to oils of petroleum origin, having utility in the processing of oils from coal, shale and other sources.

BACKGROUND OF THE INVENTION

In general, gasoline and other liquid hydrocarbon fuels boil in therange of about 100 to about 650° F. However, the crude oil from whichthese fuels are made contains a diverse mixture of hydrocarbons andother compounds which vary widely in molecular weight and therefore boilover a wide range. For example, crude oils are known in which 30 to 60%or more of the total volume of oil is composed of compounds boiling attemperatures above 650° F. Among these are crudes in which about 10% toabout 30% or more of the total volume consists of compounds so heavy inmolecular weight that they boil above 1025° F. or at least will not boilbelow 1025° F. atmospheric pressure.

Because these relatively abundant high boiling components of crude oilare unsuitable for inclusion in gasoline and other liquid hydrocarbonfuels, the petroleum refining industry has developed processes forcracking or breaking the molecules of the high molcular weight, highboiling compounds into smaller molecules which do boil over anappropriate boiling range. The cracking process which is most widelyused for this purpose is known as fluid catalytic cracking (FCC).Although the FCC process has reached a highly advanced state, and manymodified forms and variations have been developed, their unifying factoris that a vaporized hydrocarbon feedstock is caused to crack at anelevated temperature in contact with a cracking catalyst that issuspended in the feedstock vapors. Upon attainment of the desired degreeof molecular weight and boiling point reduction the catalyst isseparated from the desired products. Crude oil in the natural statecontains a variety of materials which tend to have quite troublesomeeffects on FCC processes, and only a portion of these troublesomematerials can be economically removed from the crude oil. Among thesetroublesome materials are coke precursors (such as asphaltenes,polynuclear aromatics, etc.), heavy metals (such as nickel, vanadium,iron, copper, etc.), lighter metals (such as sodium, potassium, etc.),sulfur, nitrogen and others. Certain of these, such as the lightermetals, can be economically removed by desalting operations, which arepart of the normal procedure for pretreating crude oil for fluidcatalytic cracking. Other materials, such as coke precursors,asphaltenes and the like, tend to break down into coke during thecracking operation, which coke deposits on the catalyst, impairingcontact between the hydrocarbon feedstock and the catalyst, andgenerally reducing its potency or activity level. The heavy metalstransfer almost quantitatively from the feedstock to the catalystsurface.

If the catalyst is reused again and again for processing additionalfeedstock, which is usually the case, the heavy metals can accumulate onthe catalyst to the point that they unfavorably alter the composition ofthe catalyst and/or the nature of its effect upon the feedstock. Forexample, vanadium tends to form fluxes with certain components ofcommonly used FCC catalysts, lowering the melting point of portions ofthe catalyst particles sufficiently so that they begin to sinter andbecome ineffective cracking catalysts. Accumulations of vanadium andother heavy metals, especially nickel, also "poison" the catalyst. Theytend in varying degrees to promote excessive dehydrogenation andaromatic condensation, resulting in excessive production of carbon andgases with consequent impairment of liquid fuel yield. An oil such as acrude or crude fraction or other oil that is particularly abundant innickel and/or other metals exhibiting similar behavior, while containingrelatively large quantities of coke precursors, is referred to herein asa carbo-metallic oil, and represents a particular challenge to thepetroleum refiner.

In general the coke-forming tendency or coke precursor content of an oilcan be ascertained by determining the weight percent of carbon remainingafter a sample of that oil has been pyrolyzed. The industry accepts thisvalue as a measure of the extent to which a given oil tends to formnon-catalytic coke when employed as feedstock in a catalytic cracker.Two established tests are recognized, the Conradson Carbon andRamsbottom Carbon tests, the latter being described in ASTM Test No.D524-76. In conventional FCC practice, Ramsbottom carbon values on theorder of about 0.1 to about 1.0 are regarded as indicative of acceptablefeed. The present invention is concerned with the use of hydrocarbonfeedstocks which have higher Ramsbottom carbon values and thus exhibitsubstantially greater potential for coke formation than the usual feeds.

Since the various heavy metals are not of equal catalyst poisoningactivity, it is convenient to express the poisoning activity of an oilcontaining a given poisoning metal or metals in terms of the amount of asingle metal which is estimated to have equivalent poisoning activity.Thus, the heavy metals content of an oil can be expressed by thefollowing formula (patterned after that of W. L. Nelson in Oil and GasJournal, page 143, Oct. 23, 1961) in which the content of each metalpresent is expressed in parts per million of such metal, as metal, on aweight basis, based on the weight of feed:

    Nickel Equivalents=Ni+(V/4.8)+(Fe/7.1)+(Cu/1.23)

According to conventional FCC practice, the heavy metal content offeedstock for FCC processing is controlled at a relatively low level,e.g. about 0.25 ppm Nickel Equivalents or less. The present invention isconcerned with the processing of feedstocks containing metalssubstantially in excess of this value, and which therefore have asignificantly greater potential for accumulating on and poisoningcatalyst.

The above formula can also be employed as a measure of the accumulationof heavy metals on cracking catalyst, except that the quantity of metalemployed in the formula is based on the weight of catalyst (moisturefree basis) instead of the weight of feed. In conventional FCC practice,in which a circulating inventory of catalyst is used again and again inthe processing of fresh feed, with periodic or continuing minor additionand withdrawal of fresh and spent catalyst, the metal content of thecatalyst is maintained at a level which may for example be in the rangeof about 200 to about 600 ppm Nickel Equivalents. The process of thepresent invention is concerned with the use of catalyst having asubstantially larger metals content, and which therefore has a muchgreater than normal tendency to promote dehydrogenation, aromaticcondensation, gas production or coke formation. Therefore, such highermetals accumulation is normally regarded as quite undesirable in FCCprocessing.

There has been a long standing interest in the conversion ofcarbo-metallic oils into gasoline and other liquid fuels. For example,in the 1950s it was suggested that a variety of carbometallic oils couldbe successfully converted to gasoline and other products in theHoudresid process. Turning from the FCC mode of operation, the Houdresidprocess employed catalyst particles of "granular size" (much larger thanconventional FCC catalyst particle size) in a compact gravitating bed,rather than suspending catalyst particles in feed and product vapors ina fluidized bed. The productivity of the process, compared to fluidcatalytic cracking with lighter gas oils, was low. But the Houdresidprocess did offer some advantages. It appeared that the adverse effectspreviously encountered with heavy metals in the feed were not as great abarrier in the Houdresid process as one might expect in FCC processing.The heavy metal which accumulated on or near the outer surfaces of thecatalyst particles apparently could be removed to some extent by anattrition process, which selectively removed an outer layer ofmetal-contaminated catalyst. The catalysts were very cheap, but alsorelatively inactive, highly unsuitable by today's standards. While themaximum tolerable limit of heavy metal contamination on catalyst influid catalytic cracking was then thought to be about 200 parts permillion, the Houdresid process did continue to operate satisfactorilyeven when the total nickel plus vanadium content of the catalyst hadreached 870 ppm. Moreover, it was found that the required levels ofselectivity could be maintained without withdrawing catalyst from theprocess, except to the extent that withdrawal was required by normalmechanical losses (e.g. attrition and inadvertent discharge with offgases) and by the attrition used to control metals level. Today suchattrition of catalyst to fine particulates would present an expensiveenvironmental problem, thus considerably increasing difficultiesinvolved in practicing the process.

Although the Houdresid process obviously represented a step forward indealing with the effects of metal contamination and coke formation oncatalyst performance, its productivity was limited. Because itsoperation was uneconomical, the first Houdresid unit is no longeroperating. Thus, for the 25 years which have passed since the Houdresidprocess was first introduced commercially, the art has continued itsarduous search for suitable modifications or alternatives to the FCCprocess which would permit commercially successful operation on reducedcrude and the like. During this period a number of proposals have beenmade; some have been used commercially to a certain extent.

Several proposals involve treating the heavy oil feed to remove themetal therefrom prior to cracking, such as by hydrotreating, solventextraction and complexing with Friedel-Crafts catalysts, but thesetechniques have been criticized as unjustified economically. Anotherproposal employs a combination cracking process having "dirty oil" and"clean oil" units. Still another proposal blends residual oil with gasoil and controls the quantity of residual oil in the mixture in relationto the equilibrium flash vaporization temperature at the bottom of theriser type cracker unit employed in the process. Still another proposalsubjects the feed to a mild preliminary hydrocracking or hydrotreatingoperation before it is introduced into the cracking unit. It has alsobeen suggested to contact a carbo-metallic oil such as reduced crudewith hot taconite pellets to produce gasoline. This is a small samplingof the many proposals which have appeared in the patent literature andtechnical reports.

Notwithstanding the great effort which has been expended and the factthat each of these proposals overcomes some of the difficultiesinvolved, conventional FCC practice today bears mute testimony to thedearth of carbo-metallic oil-cracking techniques that are botheconomical and highly practical in terms of technical feasibility. Somecrude oils are relatively free of coke precursors or heavy metals orboth, and the troublesome components of crude oil are for the most partconcentrated in the highest boiling fractions. Accordingly, it has beenpossible to largely avoid the problems of coke precursors and heavymetals by sacrificing the liquid fuel yield which would be potentiallyavailable from the highest boiling fractions. More particularly,conventional FCC practice has employed as feedstock that fraction ofcrude oil which boils at about 650° F. to about 1000° F., such fractionsbeing relatively free of coke precursors and heavy metal contamination.Such feedstock, known as "vacuum gas oil" (VGO) is generally preparedfrom crude oil by distilling off the fractions boiling below about 650°F. at atmospheric pressure and then separating by further vacuumdistillation from the heavier fractions a cut boiling between about 650°F. and about 900° to 1025° F.

The vacuum gas oil is used as feedstock for conventional FCC processing.The heavier fractions are normally employed for a variety of otherpurposes, such as for instance production of asphalt, residual fuel oil,#6 fuel oil, or marine Bunker C fuel oil, which represents a great wasteof the potential value of this portion of the crude oil, especially inlight of the great effort and expense which the art has been willing toexpend in the attempt to produce generally similar materials from coaland shale oils. The present invention is aimed at the simultaneouscracking of these heavier fractions containing substantial quantities ofboth coke precursors and heavy metals, and possibly other troublesomecomponents, in conjunction with the lighter oils, thereby increasing theoverall yield of gasoline and other hydrocarbon liquid fuels from agiven quantity of crude. As indicated above, the present invention by nomeans constitutes the first attempt to develop such a process, but thelong standing recognition of the desirability of cracking carbo-metallicfeedstocks, along with the slow progress of the industry toward doingso, show the continuing need for such a process. It is believed that thepresent process is uniquely advantageous for dealing with the problem oftreating such carbo-metallic oils in an economically and technicallysound manner.

SUMMARY OF THE INVENTION

The present invention is notable in providing a simple, relativelystraightforward and highly productive approach to the conversion ofcarbo-metallic feed such as reduced crude or the like to various lighterproducts such as gasoline. The carbo-metallic feed comprises or iscomposed of oil which boils above about 650° F. This oil, hereinaftersometimes referred to as converter feed, preferably contains at leastabout 70% by volume of 650° F.+ material, preferably includes at leastabout 10% by volume of material which will not boil below about 1025°F., and preferably has had substantially no prior hydrotreatment. Suchoil, or at least the 650° F.+ portion thereof, is characterized by aheavy metal content of at least about 4, preferably more than about 5,and most preferably at least about 5.5. ppm of Nickel Equivalents byweight and by a carbon residue on pyrolysis of at least about 1% andmore preferably at least about 2% by weight.

In accordance with the invention, the carbo-metallic feed, in the formof a pumpable liquid, is brought into contact with hot conversioncatalyst in a weight ratio of catalyst to feed in the range of about 3to about 18 and preferably more than about 6. The catalyst has anequilibrium microactivity test conversion activity level of at leastabout 40 and more preferably at least about 60 volume percent. Also, thecatalyst bears an accumulation of heavy metal(s) corresponding to atleast about 3000 ppm, more preferably at least about 4000 ppm and stillmore preferably at least about 5000 ppm by weight of Nickel Equivalents,expressed as metal(s) on regenerated equilibrium catalyst.

In accordance with the invention, the carbo-metallic feed is alsobrought into contact with one or more additional materials includingparticularly liquid water in a weight ratio relative to feed rangingfrom about 0.04 to about 0.15, more preferably about 0.04 to about 0.1and still more preferably about 0.05 to about 0.1. Such additionalmaterials, including the liquid water, may be brought into admixturewith the feed prior to, during or after mixing the feed with theaforementioned catalyst, and either after or, preferably, before,vaporization of the feed.

The feed, catalyst and water (e.g. in the form of liquid water or in theform of steam produced by vaporization of liquid water in contact withthe feed) are introduced at one or more points into a progressive flowtype reactor. While the mixture of feed, catalyst and steam produced byvaporization of the liquid water in contact with the feed flows throughthe reactor, the feed undergoes a conversion step which includescracking. The reactor includes an elongated reaction chamber preferablya riser which is at least partly vertical or inclined and in which thefeed material, catalyst, steam and resultant products are maintained incontact with one another while flowing as a dilute phase or stream at alineal velocity of at least about 25 and preferably at least about 35feet per second, and for a vapor residence time in the range of about0.5 to 10 seconds, preferably about 6 seconds or less and still morepreferably about 3 seconds or less.

The reaction is conducted at a temperature of about 975° to about 1300°F. and more preferably about 985° to about 1200° F., measured at thereaction chamber exit, under a total pressure of about 10 to about 50and preferably about 15 to about 35 psia (pounds per square inchabsolute) under conditions sufficiently severe to provide a conversionper pass in the range of about 60 to about 90, and still more preferablyabout 70 to about 85, volume percent and to lay down coke on thecatalyst in an amount in the range of about 0.3 to about 3% by weightand preferably at least about 0.5%. The overall rate of coke production,based on weight of fresh feed, is in the range of about 4 to about 14%by weight. The process can be operated with the foregoing conditionsbeing insufficient to fully vaporize the converter feed.

At the end of the predetermined residence time, said catalyst isseparated from at least a substantial portion of the stream comprisingsaid catalyst, steam and resultant cracking products formed in theelongated reaction chamber. After stripping and regeneration thecatalyst is returned to the reactor for contact with fresh feed.

Depending on how the process of the invention is practiced, one or moreof the following advantages may be realized. If desired, and preferably,the process may be operated without added hydrogen in the reactionchamber. If desired, and preferably, the process may be operated withoutprior hydrotreating of the feed and/or without other process of removalof asphaltenes or metals from the feed, and this is true even where thecarbo-metallic oil as a whole contains more than about 4, or more thanabout 5 or even more than about 5.5 ppm Nickel Equivalents by weight ofheavy metals and has a carbon residue on pyrolysis greater than about1%, greater than about 1.4% or greater than about 2% by weight.Moreover, all of the converter feed, as above described, may be crackedin one and the same conversion chamber. The cracking reaction may becarried out with a catalyst which has previously been used (recycled,except for such replacement as required to compensate for normal lossesand deactivation) to crack a carbo-metallic feed under the abovedescribed conditions. Heavy hydrocarbons not cracked to gasoline in afirst pass may be recycled with or without hydrotreating for furthercracking in contact with the same kind of feed in which they were firstsubjected to cracking conditions, and under the same kind of conditions;but operation in a substantially once-through or single pass mode (e.g.less than about 15% by volume of recycle based on volume of fresh feed)is preferred.

The process as above described may be practiced in conjunction with anumber of preferred alternatives, refinements or more commonlyencountered conditions, a few of which will be referred to under theheading "Description of Various and Preferred Embodiments" below.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic diagram of a first apparatus for carrying out theinvention.

FIG. 2 is a schematic diagram of a second apparatus for carrying out theinvention.

DESCRIPTION OF VARIOUS AND PREFERRED EMBODIMENTS Carbo-Metallic OilConverter Feed

The present invention provides a process for the continuous catalyticconversion of a wide variety of carbometallic oils to lower molecularweight products, while maximizing production of highly valuable liquidproducts, and making it possible, if desired, to avoid vacuumdistillation and other expensive treatments such as hydrotreating. Theterm "oils", includes not only those predominantly hydrocarboncompositions which are liquid at room temperature (i.e., 68° F.), butalso those predominantly hydrocarbon compositions which are asphalts ortars at ambient temperature but liquify when heated to temperatures inthe range of up to about 800° F. The invention is applicable tocarbo-metallic oils, whether of petroleum origin or not. For example,provided they have the requisite boiling range, carbon residue onpyrolysis and heavy metals content, the invention may be applied to theprocessing of such widely diverse materials as heavy bottoms from crudeoil, heavy bitumen crude oil, those crude oils known as "heavy crude"which approximate the properties of reduced crude, shale oil, tar sandextract, products from coal liquification and solvated coal, atmosphericand vacuum reduced crude, extracts and/or bottoms (raffinate) fromsolvent de-asphalting, aromatic extract from lube oil refining, tarbottoms, heavy cycle oil, slop oil, other refinery waste streams andmixtures of the foregoing. Such mixtures can for instance be prepared bymixing available hydrocarbon fractions, including oils, tars, pitchesand the like. Also, powdered coal may be suspended in the carbo-metallicoil. Persons skilled in the art are aware of techniques for demetalationof carbo-metallic oils, and demetalated oils may be converted using theinvention; but it is an advantage of the invention that it can employ asfeedstock carbo-metallic oils that have had no prior demetalationtreatment. Likewise, the invention can be applied to hydrotreatedfeedstocks; but it is an advantage of the invention that it cansuccessfully convert carbo-metallic oils which have had substantially noprior hydrotreatment. However, the preferred application of the processis to reduced crude, i.e., that fraction of crude oil boiling at andabove 650° F., alone or in admixture with virgin gas oils. While the useof material that has been subjected to prior vacuum distillation is notexcluded, it is an advantage of the invention that it can satisfactorilyprocess material which has had no prior vacuum distillation, thus savingon capital investment and operating costs as compared to conventionalFCC processes that require a vacuum distillation unit.

Table I below provides a comparison between a typical vacuum gas oil(VGO) which has been used heretofore in fluid catalytic cracking, withvarious reduced crudes, constituting a few examples of the many reducedcrudes useable in the present invention:

    __________________________________________________________________________                   Volume                                                                        % of 650+                                                                             Ramsbottom                                                       API  Fraction                                                                              Carbon  Wt. ppm         Wt. %                                                                              Weight of                           Grav.                                                                              Boiling at                                                                             Content                                                                              Metals.sup.(3)  S in Nitrogen                                                                              Wt.m)                       650+.sup.(2)                                                                       650-    650-                                                                              650+            Ni  650+ 650+ Fraction                                                                         ppm               Oil or Crude                                                                            Total                                                                              1025                                                                              1025+                                                                             1025                                                                              Total                                                                             Ni  V   Fe  Equiv.                                                                            Total                                                                              Total                                                                             Basic                                                                             Na                __________________________________________________________________________    VGO       28.4 100 0.0 0.38                                                                              0.38                                                                              0.2  0.1                                                                              2.6 .059                                                                               .83 722 260 0.8               Mexican Isthmus                                                                         16.9 65.3                                                                              34.7                                                                              0.49                                                                              4.96                                                                              2.5+                                                                              33.8                                                                              1.9 9.81                                                                              2.75 950 450 6.9                         (21.3)                                                              Mexican Kurkuk.sup.(1)                                                                  17.4             9.30                                                                              35.0                                                                              99.0                                                                              17.0                                                                              58.02                                                                             2.94 2100                                                                              723 1.8               Murban    23.1 78.7                                                                              21.3                                                                              0.49                                                                              3.99                                                                              3.0+                                                                               1.5                                                                              11.9                                                                              4.99                                                                              1.64 512 200 7.5               Arabian Light                                                                           19.1 64.7                                                                              35.3                                                                              0.47                                                                              6.56                                                                              6.4 24.7                                                                              3.2 12.00                                                                             2.39 940 507 9.2               Arabian Med.                                                                            14.5 51.8                                                                              48.2                                                                              0.46                                                                              9.00                                                                              19.6                                                                              63.0                                                                              2.9 33.13                                                                             4.43                           Ekofisk   22.7 72.8                                                                              27.2                                                                              0.36                                                                              4.42                                                                              1.4  3.0                                                                              2.4 2.36                                                                              0.38                           Fosterton 10.9 43.6                                                                              56.4                                                                              0.42                                                                              16.81                                                                             48.8                                                                              119.0                                                                             3.1 74.03                                                                             4.22                           Iranian Light                                                                           17.4 60.8                                                                              39.2                                                                              0.48                                                                              9.01                                                                              21.9                                                                              60.0                                                                              3.1 34.84                                                                               2.50.sup.(4)                 La./Miss Sweet                                                                          23.7 80.2                                                                              19.8                                                                              0.33                                                                              4.36                                                                              2.7+                                                                              --  8.5 3.90                                                                              0.26                           Wyoming Sour                                                                            12.4 40.7                                                                              59.3                                                                              0.32                                                                              15.1                                                                              0.6 70.0                                                                              2.0 15.47                                                                             3.84                           __________________________________________________________________________     .sup.(1) A refinery blend of Mexican and Kirkuk crudes.                       .sup.(2) Throughout the table 650 and 1025 refer to 650° F. and        1025° F. respectively; 650+ refers to 650° F.+ material as      defined below.                                                                .sup.(3) Copper level was below 0.5%, except that Mexican Kirkuk containe     0.6%; all metals expressed as metal in ppm, based on the weight of the        650+ fraction.                                                                .sup.(4) Calculated.                                                     

As can be seen from the Table, the heavier or higher boiling feeds arecharacterized by relatively lower API Gravity values than theillustrative vacuum gas oil (VGO). In general lower boiling and/orhigher API Gravity cat cracker feedstocks have been considered highlysuperior to higher boiling and/or lower API Gravity feedstocks.Comparisons of the gasoline yield of high boiling feeds compared tomedium boiling feeds at constant coke yield have shown that the mediumboiling feeds provide superior gasoline yield for a given coke yield.

In accordance with the invention one provides a carbo-metallic oilfeedstock, at least about 70%, more preferably at least about 85% andstill more preferably about 100% (by volume) of which boils at and aboveabout 650° F. All boiling temperatures herein are based on standardatmospheric pressure conditions. In carbo-metallic oil partly or whollycomposed of material which boils at and above about 650° F., suchmaterial is referred to herein as 650° F.+ material; and 650° F.+material which is part of or has been separated from an oil containingcomponents boiling above and below 650° F. may be referred to as a 650°F.+ fraction. But the terms "boils above" and "650° F.+" are notintended to imply that all of the material characterized by said termswill have the capability of boiling. The carbo-metallic oilscontemplated by the invention may contain material which may not boilunder any conditions; for example, certain asphalts and asphaltenes maycrack thermally during distillation, apparently without boiling. Thus,for example, when it is said that the feed comprises at least about 70%by volume of material which boils above about 650° F., it should beunderstood that the 70% in question may include some material which willnot boil or volatilize at any temperature. These non-boilable materials,when present, may frequently or for the most part be concentrated inportions of the feed which do not boil below about 1000° F., 1025° F. orhigher. Thus, when it is said that at least about 10%, more preferablyabout 15% and still more preferably at least about 20% (by volume) ofthe 650° F.+ fraction will not boil below about 1000° F. or 1025° F., itshould be understood that all or any part of the material not boilingbelow about 1000° or 1025° F., may or may not be volatile at and abovethe indicated temperatures.

Preferably, the contemplated feeds, or at least the 650° F.+ materialtherein, have a carbon residue on pyrolysis of at least about 2 orgreater. For example, the Ramsbottom carbon content may be in the rangeof about 2 to about 12 and most frequently at least about 4. Aparticularly common range is about 4 to about 8. Note that theillustrative VGO in Table 1 has a Ramsbottom carbon residue value of0.38, and that the 650° to 1025° F. fractions of the various reducedcrudes have Ramsbottom carbon values between about 0.3 and about 0.5,whereas the various reduced crudes as a whole (650+ Total) vary upwardsin Ramsbottom carbon value from about 4 to about 16.8, and still highervalues are contemplated.

Preferably, the feed has an average composition characterized by anatomic hydrogen to carbon ratio in the range of about 1.2 to about 1.9,and preferably about 1.3 to about 1.8.

The carbo-metallic feeds employed in accordance with the invention, orat least the 650° F.+ material therein, may contain at least about 4parts per million of Nickel Equivalents, as defined above, of which atleast about 2 parts per million is nickel (as metal, by weight).Carbo-metallic oils within the above range can be prepared from mixturesof two or more oils, some of which do and do not contain the quantitiesof Nickel Equivalents and nickel set forth above. It should also benoted that the above values for Nickel Equivalents and nickel representtime-weighted averages for a substantial period of operation of theconversion unit, such as one month, for example. It should also be notedthat the heavy metals have in certain circumstances exhibited somelessening of poisoning tendency after repeated oxidations and reductionson the catalyst, and the literature describes criteria for establishing"effective metal" values. For example, see the article by Cimbalo, etal, entitled "Deposited Metals Poison FCC Catalyst", Oil and GasJournal, May 15, 1972, pp 112-122, the contents of which areincorporated herein by reference. If considered necessary or desirable,the contents of Nickel Equivalents and nickel in the carbo-metallic oilsprocessed according to the invention may be expressed in terms of"effective metal" values. Notwithstanding the gradual reduction inpoisoning activity noted by Cimbalo, et al, the regeneration of catalystunder normal FCC regeneration conditions may not, and usually does not,severely impair the dehydrogenation, demethanation and aromaticcondensation activity of heavy metals accumulated on cracking catalyst.

It is known that about 0.2 to about 5 weight percent of "sulfur" in theform of elemental sulfur and/or its compounds (but reported as elementalsulfur based on the weight of feed) appears in FCC feeds and that thesulfur and modified forms of sulfur can find their way into theresultant gasoline product and, where lead is added, tends to reduce itssusceptibility to octane enhancement. Sulfur in the product gasolineoften requires sweetening when processing high sulfur containing crudes.To the extent that sulfur is present in the coke, it also represents apotential air pollutant since the regenerator burns it to SO₂ and SO₃.However, we have found that in our process the sulfur in the feed is onthe other hand able to inhibit heavy metal activity by maintainingmetals such as Ni, V, Cu and Fe in the sulfide form in the reactor.These sulfides are much less active than the metals themselves inpromoting dehydrogenation and coking reactions. Accordingly, it isacceptable to carry out the invention with a carbo-metallic oil havingat least about 0.3%, acceptably more than about 0.8% and more acceptablyat least about 1.5% by weight of sulfur in the 650° F.+ fraction.

The carbo-metallic oils useful in the invention may and usually docontain significant quantities of compounds containing nitrogen, asubstantial portion of which may be basic nitrogen. For example, thetotal nitrogen content of the carbo-metallic oils may be at least about0.05% by weight. Since cracking catalysts owe their cracking activity toacid sites on the catalyst surface or in its pores, basicnitrogen-containing compounds may temporarily neutralize these sites,poisoning the catalyst. However, the catalyst is not permanently damagedsince the nitrogen can be burned off the catalyst during regeneration,as a result of which the acidity of the active sites is restored.

The carbo-metallic oils may also include significant quantities ofpentane insolubles, for example at least about 0.5% by weight, and moretypically about 2% or more or even about 4% or more. These may includefor instance asphaltenes and other materials.

Alkali and alkaline earth metals generally do not tend to vaporize inlarge quantities under the distillation conditions employed indistilling crude oil to prepare the vacuum gas oils normally used as FCCfeedstocks. Rather, these metals remain for the most part in the"bottoms" fraction (the nonvaporized high boiling portion) which may forinstance be used in the production of asphalt or other by-products.However, reduced crude and other carbo-metallic oils are in many casesbottoms products, and therefore may contain significant quantities ofalkali and alkaline earth metals such as sodium. These metals depositupon the catalyst during cracking. Depending on the composition of thecatalyst and magnitude of the regeneration temperatures to which it isexposed, these metals may undergo interactions and reactions with thecatalyst (including the catalyst support) which are not normallyexperienced in processing VGO under conventional FCC processingconditions. If the catalyst characteristics and regeneration conditionsso require, one will of course take the necessary precautions to limitthe amounts of alkali and alkaline earth metal in the feed, which metalsmay enter the feed not only as brine associated with the crude oil inits natural state, but also as components of water or steam which aresupplied to the cracking unit. Thus, careful desalting of the crude usedto prepare the carbo-metallic feed may be important when the catalyst isparticularly susceptible to alkali and alkaline earth metals. In suchcircumstances, the content of such metals (hereinafter collectivelyreferred to as "sodium") in the feed can be maintained at about 1 ppm orless, based on the weight of the feedstock. Alternatively, the sodiumlevel of the feed may be keyed to that of the catalyst, so as tomaintain the sodium level of the catalyst which is in use substantiallythe same as or less than that of the replacement catalyst which ischarged to the unit.

According to a particularly preferred embodiment of the invention, thecarbo-metallic oil feedstock constitutes at least about 70% by volume ofmaterial which boils above about 650° F., and at least about 10% of thematerial which boils above about 650° F. will not boil below about 1025°F., The average composition of this 650° F.+ material may be furthercharacterized by: (a) an atomic hydrogen to carbon ratio in the range ofabout 1.3 to about 1.8: (b) a Ramsbottom carbon value of at least about2; (c) at least about four parts per million of Nickel Equivalents, asdefined above, of which at least about two parts per million is nickel(as metal, by weight); and (d) at least one of the following: (i) atleast about 0.3% by weight of sulfur, (ii) at least about 0.05% byweight of nitrogen, and (iii) at least about 0.5% by weight of pentaneinsolubles. Very commonly, the preferred feed will include all of (i),(ii), and (iii), and other components found in oils of petroleum andnon-petroleum origin may also be present in varying quantities providingthey do not prevent operation of the process.

Although there is no intention of excluding the possibility of using afeedstock which has previously been subjected to some cracking, thepresent invention has the definite advantage that it can successfullyproduce large conversions and very substantial yields of liquidhydrocarbon fuels from carbo-metallic oils which have not been subjectedto any substantial amount of cracking. Thus, for example, andpreferably, at least about 85%, more preferably at least about 90% andmost preferably substantially all of the carbo-metallic feed introducedinto the present process is oil which has not previously been contactedwith cracking catalyst under cracking conditions. Moreover, the processof the invention is suitable for operation in a substantiallyonce-through or single pass mode. Thus, the volume of recycle, if any,based on the volume of fresh feed is preferably about 15% or less andmore preferably about 10% or less.

Catalyst

In general, the weight ratio of catalyst to fresh feed (feed which hasnot previously been exposed to cracking catalyst under crackingconditions) used in the process is in the range of about 3 to about 18.Preferred and more preferred ratios are about 4 to about 12, morepreferably about 5 to about 10 and still more preferably about 6 toabout 10, a ratio of about 6 to about 8 presently being considered mostnearly optimum. Within the limitations of product quality requirements,controlling the catalyst to oil ratio at relatively low levels withinthe aforesaid ranges tends to reduce the coke yield of the process,based on fresh feed.

In conventional FCC processing of VGO, the ratio between the number ofbarrels per day of plant through-put and the total number of tons ofcatalyst undergoing circulation throughout all phases of the process canvary widely. For purposes of this disclosure, daily plant through-put isdefined as the number of barrels of fresh feed boiling above about 650°F. which that plant processes per average day of operation to liquidproducts boiling below about 430° F. For example, in one commerciallysuccessful type of FCC-VGO operation, about 8 to about 12 tons ofcatalyst are under circulation in the process per 1000 barrels per dayof plant through-put. In another commercially successful process, thisratio is in the range of about 2 to 3. While the present invention maybe practiced in the range of about 2 to about 30 and more typicallyabout 2 to about 12 tons of catalyst inventory per 1000 barrels of dailyplant through-put, it is preferred to carry out the process of thepresent invention with a very small ratio of catalyst weight to dailyplant through-put. More specifically, it is preferred to carry out theprocess of the present invention with an inventory of catalyst that issufficient to contact the feed for the desired residence time in theabove indicated catalyst to oil ratio while minimizing the amount ofcatalyst inventory, relative to plant through-put, which is undergoingcirculation or being held for treatment in other phases of the processsuch as, for example, stripping, regeneration and the like. Thus, moreparticularly, it is preferred to carry out the process of the presentinvention with about 2 to about 5 and more preferably about 2 tons ofcatalyst inventory or less per thousand barrels of daily plantthrough-put.

In the practice of the invention, catalyst may be added continuously orperiodically, such as, for example, to make up for normal losses ofcatalyst from the system. Moreover, catalyst addition may be conductedin conjunction with withdrawal of catalyst, such as, for example, tomaintain or increase the average activity level of the catalyst in theunit. For example, the rate at which virgin catalyst is added to theunit may be in the range of about 0.1 to about 3, more preferably about0.15 to about 2, and most preferably to about 0.2 to about 1.5 poundsper barrel of feed. If on the other hand equilibrium catalyst from FCCoperation is to be utilized, replacement rates as high as about 5 poundsper barrel can be practiced. Where circumstances are such that thecatalyst employed in the unit is below average in resistance todeactivation and/or conditions prevailing in the unit tend to promotemore rapid deactivation, one may employ rates of addition greater thanthose stated above; but in the opposite circumstances, lower rates ofaddition may be employed.

Without wishing to be bound by any theory, it appears that a number offeatures of the process to be described in greater detail below, suchas, for instance, the residence time and the admixture of water with thefeedstock, tend to restrict the extent to which cracking conditionsproduce metals in the reduced state on the catalyst from heavy metalsulfide(s), sulfate(s) or oxide(s) deposited on the catalyst particlesby prior exposures to carbo-metallic feedstock and regenerationconditions. Thus, the process appears to afford significant control overthe poisoning effect of heavy metals on the catalyst, even when theaccumulations of such metals are quite substantial.

Accordingly, the process may be practiced with catalyst bearingaccumulations of heavy metals which heretofore would have beenconsidered quite intolerable in conventional FCC-VGO operations. Forthese reasons, operation of the process with catalyst bearing heavymetals accumulations in the range of about 3,000 to about 70,000 ppmNickel Equivalents, on the average is contemplated. More specifically,the accumulation may be in the range of about 4,000 to about 50,000 ppmand particularly more than about 5,000 to about 30,000 ppm. Theforegoing ranges are based on parts per million of Nickel Equivalents,in which the metals are expressed as metal, by weight, measured on andbased on regenerated equilibrium catalyst. For example, one might employequilibrium catalyst from another unit, for example, an FCC unit whichhas been used in the cracking of a feed, e.g. vacuum gas oil, having acarbon residue on pyrolysis of less than 1 and containing less thanabout 4 ppm Nickel Equivalents of heavy metals.

In any event, the equilibrium concentration of heavy metals in thecirculating inventory of catalyst can be controlled (includingmaintained or varied as desired or needed) by manipulation of the rateof catalyst addition discussed above. Thus, for example, addition ofcatalyst may be maintained at a rate which will control the heavy metalsaccumulation on the catalyst in one of the ranges set forth above.

In general, it is preferred to employ a catalyst having a relativelyhigh level of cracking activity, providing high levels of conversion andproductivity at low residence times. The conversion capabilities of thecatalyst may be expressed in terms of the conversion produced duringactual operation of the process and/or in terms of conversion producedin standard catalyst activity tests. For example, it is preferred toemploy catalyst which, in the course of extended operation in theprocess, is sufficiently active for sustaining a level of conversion ofat least about 60%. In this connection, conversion is expressed inliquid volume percent, based on fresh feed. Also, for example, thepreferred catalyst may be defined as one which, in its virgin orequilibrium state, exhibits a specified activity expressed as a volumepercentage derived by the MAT (micro-activity test). For purposes of thepresent invention the foregoing percentage is the volume percentage ofstandard feedstock that is converted to 430° F. end point gasoline andlighter products at 900° F., 16 whsv (weight hourly space velocity),calculated on the basis of catalyst dried at 1100° F.) and 3C/O(catalyst to oil ratio) by the tentative ASTM MAT test developed by ASTMCommittee D-32 using an appropriate standard feedstock, e.g. DavisonWHPS-12 primary gas oil, having the following analysis and properties:

    ______________________________________                                        API Gravity at 60° F., degrees                                                             31.0                                                      Specific Gravity at 60° F., g/cc                                                           0.8708                                                    Ramsbottom Carbon, wt. %                                                                          0.09                                                      Conradson Carbon, wt. % (est.)                                                                    0.04                                                      Carbon, wt. %       84.92                                                     Hydrogen, wt. %     12.94                                                     Sulfur, wt. %       0.68                                                      Nitrogen, ppm       305                                                       Viscosity at 100° F., centistokes                                                          10.36                                                     Watson K Factor     11.93                                                     Aniline Point       182                                                       Bromine No.         2.2                                                       Paraffins, Vol. %   31.7                                                      Olefins, Vol. %     1.8                                                       Naphthenes, Vol. %  44.0                                                      Aromatics, Vol. %   22.7                                                      Average Molecular Weight                                                                          284                                                       Nickel              Trace                                                     Vanadium            Trace                                                     Iron                Trace                                                     Sodium              Trace                                                     Chlorides           Trace                                                     B S & W             Trace                                                     ______________________________________                                        Distillation, °F.                                                                          ASTM D-1160                                               ______________________________________                                        IBP                 445                                                       10%                 601                                                       30%                 664                                                       50%                 701                                                       70%                 734                                                       90%                 787                                                       FBP                 834                                                       ______________________________________                                    

The end point of the gasoline produced in the MAT test is often definedas 430° F. tbp (true boiling point) which is a standard laboratorydistillation, but other end points could serve equally well for ourpresent purposes. Conversion is calculated by subtracting from 100 thevolume percent (based on fresh feed) of those products heavier thangasoline which remain in the recovered product.

The catalyst may be introduced into the process in its virgin form or,as previously indicated, in other than virgin form; e.g. one may useequilibrium catalyst withdrawn from another unit, such as catalyst thathas been employed in the cracking of a different feed. Whencharacterized on the basis of MAT activity, the preferred catalysts maybe described on the basis of their MAT activity "as introduced" into theprocess of the present invention, or on the basis of their "aswithdrawn" or equilibrium MAT activity in the process of the presentinvention, or on both of these bases. A preferred MAT activity forvirgin and non-virgin catalyst "as introduced" into the process of thepresent invention is at least about 60%, but it will be appreciatedthat, particularly in the case of non-virgin catalysts supplied at highaddition rates, lower MAT activity levels may be acceptable. Anacceptable "as withdrawn" or equilibrium MAT activity level of catalystwhich has been used in the process of the present invention is about 40%or more and about 60% is a preferred value.

One may employ any hydrocarbon cracking catalyst having the aboveindicated conversion capabilities. A particularly preferred class ofcatalysts includes those which have pore structures into which moleculesof feed material may enter for adsorption and/or for contact with activecatalytic sites within or adjacent the pores. Various types of catalystsare available within this classification, including for example thelayered silicates, e.g. smectites. Although the most widely availablecatalysts within this classification are the well-knownzeolite-containing catalysts, non-zeolite catalysts are alsocontemplated.

The preferred zeolite-containing catalysts may include any zeolite,whether natural, semi-synthetic or synthetic, alone or in admixture withother materials which do not significantly impair the suitability of thecatalyst, provided the resultant catalyst has the activity and porestructure referred to above. For example, if the catalyst is a mixture,it may include the zeolite component associated with or dispersed in aporous refractory inorganic oxide carrier; in such case the catalyst mayfor example contain about 1% to about 60%, more preferably about 1 toabout 40% and most typically about 5 to about 25% by weight, based onthe total weight of catalyst (water free basis) of the zeolite, thebalance of the catalyst being the porous refractory inorganic oxidealone or in combination with any of the known adjuvants for promoting orsuppressing various desired and undesired reactions. For a generalexplanation of the genus of zeolite, molecular sieve catalysts useful inthe invention, attention is drawn to the disclosures of the articlesentitled "Refinery Catalysts Are a Fluid Business" and "Making CatCrackers Work on Varied Diet", appearing respectively in the July 26,1978 and Sept. 13, 1978 issues of Chemical Week magazine. Thedescriptions of the aforementioned publications are incorporated hereinby reference.

For the most part, the zeolite components of the zeolite-containingcatalysts will be those which are known to be useful in FCC crackingprocesses. In general, these are crystalline aluminosilicates, typicallymade up of tetra coordinated aluminum atoms associated through oxygenatoms with adjacent silicon atoms in the crystal structure. However, theterm "zeolite" as used in this disclosure contemplates not onlyaluminosilicates, but also substances in which the aluminum has beenparty or wholly replaced, such as for instance by gallium and/or othermetal atoms, and further includes substances in which all or part of thesilicon has been replaced, such as for instance by germanium. Titaniumand zirconium substitution may also be practiced.

Most zeolites are prepared or occur naturally in the sodium form, sothat sodium cations are associated with the electro negative sites inthe crystal structure. The sodium cations tend to make zeolites inactiveand much less stable when exposed to hydrocarbon conversion conditions,particularly high temperatures. Accordingly, the zeolite may be ionexchanged, and where the zeolite is a component of a catalystcomposition, such ion exchanging may occur before or after incorporationof the zeolite as a component of the composition. Suitable cations forreplacement of sodium in the zeolite crystal structure include ammonium(decomposable to hydrogen), hydrogen, rare earth metals, alkaline earthmetals, etc. Various suitable ion exchange procedures and cations whichmay be exchanged into the zeolite crystal structure are well known tothose skilled in the art.

Examples of the naturally occuring crystalline aluminosilicate zeoliteswhich may be used as or included in the catalyst for the presentinvention ae faujasite, mordenite, clinoptilote, chabazite, analcite,erionite, as well as levynite, dachiardite, paulingite, noselite,ferriorite, heulandite, scolccite, stibite, harmotome, phillipsite,brewsterite, flarite, datolite, gmelinite, caumnite, leucite, lazurite,scaplite, mesolite, ptholite, nepheline, matrolite, offretite andsodalite.

Examples of the synthetic crystalline aluminosilicate zeolites which areuseful as or in the catalyst for carrying out the present invention areZeolite X, U.S. Pat. No. 2,882,244, Zeolite Y, U.S. Pat. No. 3,130,007;and Zeolite A, U.S. Pat. No. 2,882,243; as well as Zeolite B, U.S. Pat.No. 3,008,803; Zeolite D, Canada Pat. No. 661,981; Zeolite E, CanadaPat. No. 614,495; Zeolite F, U.S. Pat. No. 2,996,358; Zeolite H, U.S.Pat. No. 3,010,789; Zeolite J. U.S. Pat. No. 3,011,869; Zeolite L,Belgian Pat. No. 575,177; Zeolite M. U.S. Pat. No. 2,995,423, Zeolite O,U.S. Pat. No. 3,140,252; Zeolite Q, U.S. Pat. No. 2,991,151; Zeolite S,U.S. Pat. No. 3,054,657, Zeolite T, U.S. Pat. No. 2,950,952; Zeolite W,U.S. Pat. No. 3,012,853; Zeolite Z, Canada Pat. No. 614,495; and ZeoliteOmega, Canada Pat. No. 817,915. Also, ZK-4HJ, alpha beta and ZSM-typezeolites are useful. Moreover, the zeolites described in U.S. Pat. Nos.3,140,249, 3,140,253, 3,944,482 and 4,137,151 are also useful, thedisclosures of said patents being incorporated herein by reference.

The crystalline aluminosilicate zeolites having a faujasite-type crystalstructure are particularly preferred for use in the present invention.This includes particularly natural faujasite and Zeolite X and ZeoliteY. The crystalline aluminosilicate zeolites, such as syntheticfaujasite, will under normal conditions crystallize as regularly shaped,discrete particles of about one to about ten microns in size, and,accordingly, this is the size range frequently found in commercialcatalysts which can be used in the invention. Preferably, the particlesize of the zeolites is from about 0.5 to about 10 microns and morepreferably is from about 0.1 to about 2 microns or less. For example,zeolites prepared in situ from calcined kaolin may be characterized byeven smaller crystallites. Crystalline zeolites exhibit both an interiorand an exterior surface area, which we have defined as "portal" surfacearea, with the largest portion of the total surface area being internal.By portal surface area, we refer to the outer surface of the zeolitecrystal through which reactants are considered to pass in order toconvert to lower boiling products. Blockage of the internal channels by,for example, coke formation, blockage of entrance to the internalchannels by deposition of coke in the portal surface area, andcontamination by metals poisoning, will greatly reduce the total zeolitesurface area. Therefore, to minimize the effect of contamination andpore blockage, crystals larger than the normal size cited above arepreferably not used in the catalysts of this invention.

Commercial zeolite-containing catalysts are available with carrierscontaining a variety of metal oxides and combination thereof, includingfor example silica, alumina, magnesia, and mixtures thereof and mixturesof such oxides with clays as e.g. described in U.S. Pat. No. 3,034,948.One may for example select any of the zeolite-containing molecular sievefluid cracking catalysts which are suitable for production of gasolinefrom vacuum gas oils. However, certain advantages may be attained byjudicious selection of catalysts having marked resistance to metals. Ametal resistant zeolite catalyst is, for instance, described in U.S.Pat. No. 3,944,482, in which the catalyst contains 1-40 weight percentof a rare earth-exchanged zeolite, the balance being a refractory metaloxide having specified pore volume and size distribution. Othercatalysts described as "metals-tolerant" are described in the abovementioned Cimbalo et al article.

In general, it is preferred to employ catalysts having an over-allparticle size in the range of about 5 to about 160, more preferablyabout 40 to about 120, and most preferably about 40 to about 80 microns.

The catalyst composition may also include one or more combustionpromoters which are useful in the subsequent step of regenerating thecatalyst. Cracking of carbo-metallic oils results in substantialdeposition of coke on the catalyst, which coke reduces the activity ofthe catalyst. Thus, in order to restore the activity of the catalyst thecoke is burned off in a regeneration step, in which the coke isconverted to combustion gases including carbon monoxide and/or carbondioxide. Various substances are known which, when incorporated incracking catalyst in small quantities, tend to promote conversion of thecoke to carbon monoxide and/or carbon dioxide. Promoters of combustionto carbon monoxide tend to lower the temperature at which a given degreeof coke removal can be attained, thus diminishing the potential forthermal deactivation of the catalyst. Such promoters, normally used ineffective amounts ranging from a trace up to about 10 or 20% by weightof the catalyst, may be of any type which generally promotes combustionof carbon under regenerating conditions, or may be somewhat selective inrespect to completing the combustion of CO, or, more preferably, forreasons explained in greater detail below, may have some tendency tocombust carbon to carbon monoxide in preference to carbon dioxide.

Although a wide variety of other catalysts, including bothzeolite-containing and non-zeolite-containing may be employed in thepractice of the invention the following are examples of commerciallyavailable catalysts which have been employed in practicing theinvention:

                  TABLE II                                                        ______________________________________                                               Spe- Weight Percent                                                           cific                                                                              Zeo-                                                                     Sur- lite                                                                     face Con-                                                                     m.sup.2 /g                                                                         tent   Al.sub.2 O.sub.3                                                                      SiO.sub.2                                                                          Na.sub.2 O                                                                          Fe.sub.2 O                                                                          TiO.sub.2                         ______________________________________                                        AGZ-290  300    11.0   29.5  59.0 0.40  0.11  0.59                            GRZ-1    162    14.0   23.4  69.0 0.10  0.4   0.9                             CCZ-200  129    11.0   34.6  60.0 0.60  0.57  1.9                             Super DX 155    13.0   31.0  65.0 0.80  0.57  1.6                             F-87     240    10.0   44.0  50.0 0.80  0.70  1.6                             FOC-90   240    8.0    44.0  52.0 0.65  0.65  1.1                             HFZ 20   310    20.0   59.0  40.0 0.47  0.54  2.75                            HEZ-55   210    19.0   59.0  35.2 0.60  0.60  2.5                             ______________________________________                                    

The AGZ-290, GRZ-1, CCZ-220 and Super DX catalysts referred to above areproducts of W. R. Grace and Co. F-87 and FOC-90 are products of Filtrol,while HFZ-20 and HEZ-55 are products of Engelhard/Houdry. The above areproperties of virgin catalyst and, except in the case of zeolitecontent, are adjusted to a water free basis, i.e. based on materialignited at 1750° F. The zeolite content is derived by comparison of theX-ray intensities of a catalyst sample and of a standard materialcomposed of high purity sodium Y zeolite in accordance with draft #6,dated Jan. 9, 1978, of proposed ASTM Standard Method entitled"Determination of the Faujasite Content of a Catalyst." It is consideredan advantage that the process of the present invention can be conductedin the substantial absence of tin and/or antimony or at least in thepresence of a catalyst which is substantially free of either or both ofthese metals.

Additional Materials

The process of the present invention may be operated with the abovedescribed carbo-metallic oil and catalyst, and with H₂ O assubstantially the only additional material charged to the reaction zone.But the charging of other additional materials is not excluded. Thecharging of recycled oil to the reaction zone has already beenmentioned. As described in greater detail below, still other additionalmaterials fulfilling a variety of functions may also be charged.

In general the H₂ O and other additional materials which may be usedeach perform one or more of the following functions which offersignificant advantages over the process as performed with only thecarbo-metallic oil and catalyst. Among these functions are: controllingthe effects of heavy metals and other catalyst contaminants; enhancingcatalyst activity; absorbing excess heat in the catalyst as receivedfrom the regenerator; disposal of pollutants or conversion thereof to aform or forms in which they may be more readily separated from productsand/or disposed of; controlling catalyst temperature; diluting thecarbo-metallic oil vapors to reduce their partial pressure and increasethe yield of desired products; adjusting feed/catalyst contact time;donation of hydrogen to a hydrogen deficient carbo-metallic oilfeedstock; assisting in the dispersion of the feed; and possibly alsodistillation of products. Certain of the metals in the heavy metalsaccumulation on the catalyst are more active in promoting undesiredreactions when they are in the form of elemental metal, than they arewhen in the oxidized form produced by contact with oxygen in thecatalyst regenerator. However, the time of contact between catalyst andvapors of feed and product in past conventional catalytic cracking wassufficient so that hydrogen released in the cracking reaction was ableto reconvert a significant portion of the less harmful oxides back tothe more harmful elemental heavy metals. One can take advantage of thissituation through the introduction of additional materials which are ingaseous (including vaporous) form in the reaction zone in admixture withthe catalyst and vapors of feed and products. The increased volume ofmaterial in the reaction zone resulting from the presence of suchadditional materials tends to increase the velocity of flow through thereaction zone with a corresponding decrease in the residence time of thecatalyst and oxidized heavy metals borne thereby. Because of thisreduced residence time, there is less opportunity for reduction of theoxidized heavy metals to elemental form and therefore less of theharmful elemental metals are available for contacting the feed andproducts.

Added materials may be introduced into the process in any suitablefashion, some examples of which follow. For instance, they may beadmixed with the carbo-metallic oil feedstock prior to contact of thelatter with the catalyst. Alternatively, the added materials may, ifdesired, be admixed with the catalyst prior to contact of the latterwith the feedstock. Separate portions of the added materials may beseparately admixed with both catalyst and carbo-metallic oil. Moreover,the feedstock, catalyst and additional materials may, if desired, bebrought together substantially simultaneously. A portion of the addedmaterials may be mixed with catalyst and/or carbo-metallic oil in any ofthe above described ways, while additional portions are subsequentlybrought into admixture. For example, a portion of the added materialsmay be added to the carbo-metallic oil and/or to the catalyst beforethey reach the reaction zone, while another portion of the addedmaterials is introduced directly into the reaction zone. The addedmaterials may be introduced at a plurality of spaced locations in thereaction zone or along the length thereof, if elongated.

The amount of additional materials which may be present in the feed,catalyst or reaction zone for carrying out the above functions, andothers, may be varied as desired; but said amount will preferably besufficient to substantially heat balance the process. These materialsmay for example be introduced into the reaction zone in a weight ratiorelative to feed of up to about 0.4, preferably in the range of about0.02 to about 0.4, more preferably about 0.03 to about 0.3 and mostpreferably about 0.05 to about 0.25.

Many or all of the above desireable functions may be attained byintroducing H₂ O to the reaction zone in the form of liquid water aloneor in combination with steam. Without wishing to be bound by any theory,it appears that the use of H₂ O tends to inhibit reduction ofcatalyst-borne oxides, sulfites and sulfides to the free metallic formwhich is believed to promote condensation-dehydrogenation withconsequent promotion of coke and hydrogen yield and accompanying loss ofproduct. Moreover, H₂ O may also, to some extent, reduce deposition ofmetals onto the catalyst surface. There may also be some tendency todesorb nitrogen-containing and other heavy contaminant-containingmolecules from the surface of the catalyst particles, or at least sometendency to inhibit their absorption by the catalyst. It is alsobelieved that added H₂ O tends to increase the acidity of the catalystby Bronsted acid formation which in turn enhances the activity of thecatalyst. Assuming the H₂ O as supplied is cooler than the regeneratedcatalyst and/or the temperature of the reaction zone, the sensible heatinvolved in raising the temperature of the H₂ O upon contacting thecatalyst in the reaction zone or elsewhere can absorb excess heat fromthe catalyst. Where the H₂ O is or includes recycled water that containsfor example about 500 to about 5000 ppm of H₂ S dissolved therein, anumber of additional advantages may accrue. The ecologicallyunattractive H₂ S need not be vented to the atmosphere, the recycledwater does not require further treatment to remove H₂ S and the H₂ S maybe of assistance in reducing coking of the catalyst by passivation ofthe heavy metals, i.e. by conversion thereof to the sulfide form whichhas a lesser tendency than the free metals to enhance coke and hydrogenproduction. In the reaction zone, the presence of H₂ O can dilute thecarbo-metallic oil vapors, thus reducing their partial pressure andtending to increase the yield of the desired products. It has beenreported that H₂ O is useful in combination with other materials ingenerating hydrogen during cracking; thus it may be able to act as ahydrogen donor for hydrogen deficient carbo-metallic oil feedstocks. TheH₂ O may also serve certain purely mechanical functions such as:assisting in the atomizing or dispersion of the feed; competing withhigh molecular weight molecules for adsorption on the surface of thecatalyst, thus interrupting coke formation; steam distillation ofvaporizable product from unvaporized feed material; and disengagement ofproduct from catalyst upon conclusion of the cracking reaction. It isparticularly preferred to bring together H₂ O, catalyst andcarbo-metallic oil substantially simultaneously. For example, one mayadmix H₂ O and feedstock in an atomizing nozzle and immediately directthe resultant spray into contact with the catalyst at the downstream endof the reaction zone.

The addition of steam to the reaction zone is frequently mentioned inthe literature of fluid catalytic cracking. Addition of liquid water tothe feed is discussed relatively infrequently, compared to theintroduction of steam directly into the reaction zone. However, inaccordance with the present invention it is particularly preferred thatliquid water be brought into intimate admixture with the carbo-metallicoil in a weight ratio of about 0.04 to about 0.15 at or prior to thetime of introduction of the oil into the reaction zone, whereby thewater (e.g., in the form of liquid water or in the form of steamproduced by vaporization of liquid water in contact with the oil) entersthe reaction zone as part of the flow of feedstock which enters suchzone. Although not wishing to be bound by any theory, it is believedthat the foregoing is advantageous in promoting dispersion of thefeedstock. Also, the heat of vaporization of the water, which heat isabsorbed from the catalyst, from the feedstock, or from both, causes thewater to be a more efficient heat sink than steam alone. Preferably theweight ratio of liquid water to feed is about 0.04 to about 0.1, morepreferably about 0.05 to about 0.1.

Of course, the liquid water may be introduced into the process in theabove described manner or in other ways, and in either event theintroduction of liquid water may be accompanied by the introduction ofadditional amounts of water as steam into the same or different portionsof the reaction zone or into the catalyst and/or feedstock. For example,the amount of additional steam may be in a weight ratio relative to feedin the range of about 0.01 to about 0.25, with the weight ratio of totalH₂ O (as steam and liquid water) to feedstock being about 0.3 or less.The charging weight ratio of liquid water relative to steam in suchcombined use of liquid water and steam may thus range from about 5 toabout 0.2. Such ratio may be maintained at a predetermined level withinsuch range or varied as necessary or desired to adjust or maintain theheat balance of the reaction.

Other materials may be added to the reaction zone to perform one or moreof the above described functions. For example, thedehydrogenation-condensation activity of heavy metals may be inhibitedby introducing hydrogen sulfide gas into the reaction zone. Hydrogen maybe made available for hydrogen deficient carbo-metallic oil feedstocksby introducing into the reaction zone either a conventional hydrogendonor diluent such as a heavy naphtha or relatively low molecular weightcarbon-hydrogen fragment contributors, including for example: lightparaffins; low molecular weight alcohols and other compounds whichpermit or favor intermolecular hydrogen transfer; and compounds thatchemically combine to generate hydrogen in the reaction zone such as byreaction of carbon monoxide with water, or with alcohols, or witholefins, or with other materials or mixtures of the foregoing.

All of the above mentioned additional materials (including water), aloneor in conjunction with each other or in conjunction with othermaterials, such as nitrogen or other inert gases, light hydrocarbons,and others, may perform any of the above-described functions for whichthey are suitable, including without limitation, acting as diluents toreduce feed partial pressure and/or as heat sinks to absorb excess heatpresent in the catalyst as received from the regeneration step. Theforegoing is a discussion of some of the functions which can beperformed by materials other than catalyst and carbo-metallic oilfeedstock introduced into the reaction zone, and it should be understoodthat other materials may be added or other functions performed withoutdeparting from the spirit of the invention.

Illustrative Apparatus

The invention may be practiced in a wide variety of apparatus. However,the preferred apparatus includes means for rapidly vaporizing as muchfeed as possible and efficiently admixing feed, water and catalyst(although not necessarily in that order), for causing the resultantmixture to flow as a dilute suspension in a progressive flow mode, andfor separating the catalyst from cracked products and any uncracked oronly partially cracked feed at the end of a predetermined residence timeor times, it being preferred that all or at least a substantial portionof the product should be abruptly separated from at least a portion ofthe catalyst.

For example, the apparatus may include, along its elongated reactionchamber, one or more points for introduction of carbo-metallic feed, oneor more points for introduction of catalyst, one or more points forintroduction of additional materials including water, one or more pointsfor withdrawal of products and one or more points for withdrawal ofcatalyst. The means for introducing feed, catalyst and other materialmay range from open pipes to sophisticated jets or spray nozzles, itbeing preferred to use means capable of breaking up the liquid feed intofine droplets. Preferably, the catalyst, liquid water (when used) andfresh feed are brought together in an apparatus similar to thatdisclosed in U.S. Patent Application Ser. No. 969,601 of George D. Myerset al, filed Dec. 14, 1978, the entire disclosure of which is herebyincorporated herein by reference. It is preferred that the reactionchamber, or at least the major portion thereof, be more nearly verticalthan horizontal and have a length to diameter ratio of at least about10, more preferably about 20 or 25 or more. Use of a vertical riser typereactor is preferred. If tubular, the reactor can be of uniform diameterthroughout or may be provided with a continuous or step-wise increase indiameter along the reaction path to maintain or vary the velocity alongthe flow path.

In general, the charging means (for catalyst, water and feed) and thereactor configuration are such as to provide a relatively high velocityof flow and dilute suspension of catalyst. For example, the vapor orcatalyst velocity in the riser will be usually at least about 25 andmore typically at least about 35 feet per second. This velocity mayrange up to about 55 or about 75 feet per second or higher. The velocitycapabilities of the reactor will in general be sufficient to preventsubstantial build-up of a catalyst bed in the bottom or other portionsof the riser, whereby the catalyst loading in the riser can bemaintained below about 4 or 5 pounds and below about 2 pounds per cubicfoot, respectively, at the upstream (e.g. bottom) and downstream (e.g.top) ends of the riser.

The progressive flow mode involves, for example, flowing of catalyst,feed and steam as a stream in a positively controlled and maintaineddirection established by the elongated nature of the reaction zone. Thisis not to suggest however that there must be strictly linear flow. As iswell known, turbulent flow and "slippage" of catalyst may occur to someextent especially in certain ranges of vapor velocity and some catalystloadings, although it has been reported adviseable to employsufficiently low catalyst loadings to restrict slippage and back-mixing.

Most preferably the reactor is one which abruptly separates asubstantial portion or all of the vaporized cracked products from thecatalyst at one or more points along the riser, and preferably separatessubstantially all of the vaporized cracked products from the catalyst atthe downstream end of the riser. A preferred type of reactor embodiesballistic separation of catalyst and products; that is, catalyst isprojected in a direction established by the riser tube, and is caused tocontinue its motion in the general direction so established, while theproducts, having lesser momentum, are caused to make an abrupt change ofdirection, resulting in an abrupt, substantially instantaneousseparation of product from catalyst. In a preferred embodiment referredto as a vented riser, the riser tube is provided with a substantiallyunobstructed discharge opening at its downstream end for discharge ofcatalyst. An exit port in the side of the tube adjacent the downstreamend receives the products. The discharge opening communicates with acatalyst flow path which extends to the usual stripper and regenerator,while the exit port communicates with a product flow path which issubstantially or entirely separated from the catalyst flow path andleads to separation means for separating the products from therelatively small portion of catalyst, if any, which manages to gainentry to the product exit port. Examples of a ballistic separationapparatus and technique as above described, are found in U.S. Pat. Nos.4,066,533 and 4,070,159 to Myers et al, the disclosures of which patentsare hereby incorporated herein by reference in their entireties.

PREFERRED OPERATING CONDITIONS

Preferred conditions for operation of the process are described below.Among these are feed, catalyst and reaction temperatures, reaction andfeed pressures, residence time and levels of conversion, coke productionand coke laydown on catalyst.

In conventional FCC operations with VGO, the feedstock is customarilypreheated, often to temperatures significantly higher than are requiredto make the feed sufficiently fluid for pumping and for introductioninto the reactor. For example, preheat temperatures as high as about700° or 800° F. have been reported. But in our process as presentlypracticed it is preferred to restrict preheating of the feed, so thatthe feed is capable of absorbing a larger amount of heat from thecatalyst while the catalyst raises the feed to conversion temperature,at the same time minimizing utilization of external fuels to heat thefeedstock. Thus, where the nature of the feedstock permits, it may befed at ambient temperature. Heavier stocks may be fed at preheattemperatures of up to about 600° F., typically about 200° F. to about500° F., but higher preheat temperatures are not necessarily excluded.

The catalyst fed to the reactor may vary widely in temperature, forexample from about 1100° to about 1600° F., more preferably about 1200°to about 1500° F. and most preferably about 1300° to about 1400° F.,with about 1325° to about 1375° being considered optimum at present.

As indicated previously, the conversion of the carbo-metallic oil tolower molecular weight products may be conducted at a temperature ofabout 975° to about 1300° F., measured at the reaction chamber outlet.The reaction temperature as measured at said outlet is more preferablymaintained in the range of about 985° to about 1200° F., and mostpreferably about 1000° to about 1150° F. Depending upon the temperatureselected and the properties of the feed, all of the feed may or may notvaporize in the riser.

Although the pressure in the reactor may, as indicated above, range fromabout 10 to about 50 psia, preferred and more preferred pressure rangesare about 15 to about 35 and about 20 to about 35. In general, thepartial (or total) pressure of the feed may be in the range of about 3to about 30, more preferably about 7 to about 25 and most preferablyabout 10 to about 17 psia. The feed partial pressure may be controlledor suppressed by the introduction of gaseous (including vaporous)materials into the reactor, such as for instance the steam, water, andother additional materials described above. The process has for examplebeen operated with the ratio of feed partial pressure relative to totalpressure in the riser in the range of about 0.2 to about 0.8, moretypically about 0.3 to about 0.7 and still more typically about 0.4 toabout 0.6, with the ratio of the partial pressure of added gaseousmaterial (which includes the steam resulting from introduction of H₂ Oto the riser and may also include recycled gases) relative to totalpressure in the riser correspondingly ranging from about 0.8 to about0.2, more typically about 0.7 to about 0.3 and still more typicallyabout 0.6 to about 0.4. In the illustrative operations just described,the ratio of the partial pressure of the added gaseous material relativeto the partial pressure of the feed has been in the range of about 0.25to about 4, more typically about 0.4 to about 2.3 and still moretypically about 0.7 to about 1.7.

Although the residence time of feed and product vapors in the riser maybe in the range of about 0.5 to about 10 seconds, as described above,preferred and more preferred values are about 0.5 to about 6 and about 1to about 4 seconds, with about 1.5 to about 3.0 seconds currently beingconsidered about optimum. For example, the process has been operatedwith a riser vapor residence time of about 2.5 seconds or less byintroduction of copious amounts of gaseous materials into the riser,such amounts being sufficient to provide for example a partial pressureratio of added gaseous materials relative to hydrocarbon feed of about0.8 or more. By way of further illustration, the process has beenoperated with said residence time being about two seconds or less, withthe aforesaid ratio being in the range of about 1 to about 2. Thecombination of low feed partial pressure, very low residence time andballistic separation of products from catalyst are considered especiallybeneficial for the conversion of carbon-metallic oils. Additionalbenefits may be obtained in the foregoing combination when there is asubstantial partial pressure of added gaseous material, especially H₂ O,as described above.

Depending upon whether there is slippage between the catalyst andhydrocarbon vapors in the riser, the catalyst riser residence time mayor may not be the same as that of the vapors. Thus, the ratio of averagecatalyst reactor residence time versus vapor reactor residence time,i.e. slippage, may be in the range of about 1 to about 5, morepreferably about 1 to about 4 and most preferably about 1.2 to about 3,with about 1.2 to about 2 currently being considered optimum.

In certain types of known FCC units, there is a riser which dischargescatalyst and product vapors together into an enlarged chamber, usuallyconsidered to be part of the reactor, in which the catalyst isdisengaged from product and collected. Continued contact of catalyst,uncracked feed (if any) and cracked products in such enlarged chamberresults in an overall catalyst feed contact time appreciably exceedingthe riser tube residence times of the vapors and catalysts. Whenpracticing the process of the present invention with ballisticseparation of catalyst and vapors at the downstream (e.g. upper)extremity of the riser, such as is taught in the above mentioned Myerset al patents, the riser residence time and the catalyst contact timeare substantially the same for a major portion of the feed and productvapors. It is considered advantageous if the vapor riser residence timeand vapor catalyst contact time are substantially the same for at leastabout 80%, more preferably at least about 90% and most preferably atleast about 95% by volume of the total feed and product vapors passingthrough the riser. By denying such vapors continued contact withcatalyst in a catalyst disengagement and collection chamber one mayavoid a tendency toward re-cracking and diminished selectivity.

In general, the combination of catalyst to oil ratio, temperatures,pressures and residence times should be such as to effect a substantialconversion of the carbo-metallic oil feedstock. It is an advantage ofthe process that very high levels of conversion can be attained in asingle pass; for example the conversion may be in excess of 60% and mayrange to about 90% or higher. Preferably, the aforementioned conditionsare maintained at levels sufficient to maintain conversion levels in therange of about 60 to about 90% and more preferably about 70 to about85%. The foregoing conversion levels are calculated by subtracting from100% the percentage obtained by dividing the liquid volume of fresh feedinto 100 times the volume of liquid product boiling at and above 430° F.(tbp, standard atmospheric pressure).

These substantial levels of conversion may and usually do result inrelatively large yields of coke, such as for example about 4 to about14% by weight based on fresh feed, more commonly about 6 to about 12%and most frequently about 6 to about 10%. The coke yield can more orless quantitatively deposit upon the catalyst. At contemplated catalyststo oil ratios, the resultant coke laydown may be in excess of about 0.3,more commonly in excess of about 0.5 and very frequently in excess ofabout 1% of coke by weight, based on the weight of moisture freeregenerated catalyst. Such coke laydown may range as high as about 2%,or about 3%, or even higher.

In common with conventional FCC operations on VGO, the present processincludes stripping of spent catalyst after disengagement of the catalystfrom product vapors. Persons skilled in the art are acquainted withappropriate stripping agents and conditions for stripping spentcatalyst, but in some cases the present process may require somewhatmore severe conditions than are commonly employed. This may result, forexample, from the use of a carbo-metallic oil having constitutents whichdo not volatilize under the conditions prevailing in the reactor, whichconstituents deposit themselves at least in part on the catalyst. Suchadsorbed, unvaporized material can be troublesome from at least twostandpoints. First, if the gases (including vapors) used to strip thecatalyst can gain admission to a catalyst disengagement or collectionchamber connected to the downstream end of the riser, and if there is anaccumulation of catalyst in such chamber, vaporization of theseunvaporized hydrocarbons in the stripper can be followed by adsorptionof the bed of catalyst in the chamber. More particularly, as thecatalyst in the stripper is stripped of adsorbed feed material, theresultant feed material vapors pass through the bed of catalystaccumulated in the catalyst collection and/or disengagement chamber andmay deposit coke and/or condensed material on the catalyst in said bed.As the catalyst bearing such deposits moves from the bed and into thestripper and from thence to the regenerator, the condensed products cancreate a demand for more stripping capacity, while the coke can tend toincrease regeneration temperatures and/or demand greater regenerationcapacity. For the foregoing reasons, it is preferred to prevent orrestrict contact between stripping vapors and catalyst accumulations inthe catalyst disengagement or collection chamber. This may be done forexample by preventing such accumulations from forming, e.g. with theexception of a quantity of catalyst which essentially drops out ofcirculation and may remain at the bottom of the disengagement and/orcollection chamber, the catalyst that is in circulation may be removedfrom said chamber promptly upon settling to the bottom of the chamber.Also, to minimize regeneration temperatures and demand for regenerationcapacity, it may be desirable to employ conditions of time, temperatureand atmosphere in the stripper which are sufficient to reducepotentially volatile hydrocarbon material borne by the stripped catalystto about 10% or less by weight of the total carbon loading on thecatalyst. Such stripping may for example include reheating of thecatalyst, extensive stripping with steam, the use of gases having atemperature considered higher than normal for FCC/VGO operations, suchas for instance flue gas from the regenerator, as well as other refinerystream gases such as hydrotreater off-gas (H₂ S containing), hydrogenand others. For example, the stripper may be operated at a temperatureof about 1025° F. or higher.

Substantial conversion of carbo-metallic oils to lighter products inaccordance with the invention tends to produce sufficiently large cokeyields and coke laydown on catalyst to require some care in catalystregeneration. In order to maintain adequate activity in zeolite andnon-zeolite catalysts, it is desirable to regenerate the catalyst underconditions of time, temperature and atmosphere sufficient to reduce thepercent by weight of carbon remaining on the catalyst to about 0.25% orless, whether the catalyst bears a large heavy metals accumulation ornot. Preferably this weight precentage is about 0.1% or less and morepreferably about 0.05% or less, especially with zeolite catalysts. Theamounts of coke which must therefore be burned off of the catalysts whenprocessing carbometallic oils are usually substantially greater thanwould be the case when cracking VGO. The term coke when used to describethe present invention, should be understood to include any residualunvaporized feed or cracking product, if any such material is present onthe catalyst after stripping.

Regeneration of catalyst, burning away of coke deposited on the catalystduring the conversion of the feed, may be performed at any suitabletemperature in the range of about 1100° to about 1600° F., measured atthe regenerator catalyst outlet. This temperature is preferably in therange of about 1200° to about 1500° F., more preferably about 1275° toabout 1425° F. and optimally about 1325° to about 1375° F. The processhas been operated, for example, with a fluidized regenerator with thetemperature of the catalyst dense phase in the range of about 1300° toabout 1400° F.

When regenerating catalyst to very low levels of carbon on regeneratedcatalyst, e.g. about 0.1% or less or about 0.05% or less, based on theweight of regenerated catalyst, it is acceptable to burn off at leastabout the last 10% or at least about the last 5% by weight of coke(based on the total weight of coke on the catalyst immediately prior toregeneration) in contact with combustion producing gases containingexcess oxygen. In this connection it is contemplated that some selectedportion of the coke, ranging from all of the coke down to about the last5 or 10% by weight, can be burned with excess oxygen. By excess oxygenis meant an amount in excess of the stoichiometric requirement forburning all of the hydrogen, all of the carbon and all of the othercombustible components, if any, which are present in the above-mentionedselected portion of the coke immediately prior to regeneration. Thegaseous products of combustion conducted in the presence of excessoxygen will normally include an appreciable amount of free oxygen. Suchfree oxygen, unless removed from the by-product gases or converted tosome other form by a means or process other than regeneration, willnormally manifest itself as free oxygen in the flue gas from theregeneration unit. In order to provide sufficient driving force tocomplete the combustion of coke to low levels, when burning all or amajor portion of the coke with excess oxygen, the amount of free oxygenwill normally be not merely appreciable but substantial, i.e. there willbe a concentration of at least about 2 mole percent of free oxygen inthe total regeneration flue gas recovered from the entire, completedregeneration operation. While such technique is effective in attainingthe desired low levels of carbon on regenerated catalyst, it has itslimitations and difficulties as will become apparent from the discussionbelow.

As conventionally practiced, the burning of coke during regenerationproduces some H₂ O because of the small amount of hydrogen normallyfound in coke; but carbon monoxide and carbon dioxide are generallyregarded as the principal products. The conversion of the carbon contentof coke to carbon monoxide and carbon dioxide are highly exothermicreactions. For instance the reaction of oxygen with coke to producecarbon dioxide produces 14,108 BTUs per pound of coke, while thereaction of oxygen with coke or carbon to form carbon monoxide producesapproximately 3967 BTUs per pound of coke. The larger the amount of cokewhich must be burned from a given weight of catalyst, the greater theamount of heat released during combustion in the regenerator.

Heat released by combustion of coke in the regenerator is absorbed bythe catalyst and can be readily retained thereby until the regeneratedcatalyst is brought into contact with fresh feed. When processingcarbo-metallic oils to the relatively high levels of conversion involvedin the present invention, the amount of regenerator heat which istransmitted to fresh feed by way of recycling regenerated catalyst cansubstantially exceed the level of heat input which is appropriate in theriser for heating and vaporizing the feed and other materials, forsupplying the endothermic heat of reaction for cracking, for making upthe heat losses of the unit and so forth. Thus, in accordance with theinvention, the amount of regenerator heat transmitted to fresh feed maybe controlled, or restricted where necessary, within certain approximateranges. The amount of heat so transmitted may for example be in therange of about 500 to about 1200, more particularly about 600 to about900, and more particularly about 650 to about 850 BTUs per pound offresh feed. The aforesaid ranges refer to the combined heat, in BTUs perpound of fresh feed, which is transmitted by the catalyst to the feedand reaction products (between the contacting of feed with catalyst andthe separation of product from catalyst) for supplying the heat ofreaction (e.g. for cracking) and the difference in enthalpy between theproducts and the fresh feed. Not included in the foregoing are the heatmade available in the reactor by the adsorption of coke on the catalyst,nor the heat consumed by heating, vaporizing or reacting recycle streamsand such added materials as water, steam, naphtha and other hydrogendonors, flue gases and inert gases, or by radiation and other losses.

One or a combination of techniques may be utilized in this invention forcontrolling or restricting the amount of regeneration heat transmittedvia catalyst to fresh feed. For example, one may add a combustionpromotor to the cracking catalyst in order to reduce the temperature ofcombustion of coke to carbon dioxide and/or carbon monoxide in theregenerator. Moreover, one may remove heat from the catalyst throughheat exchange means, including for example heat exchangers (e.g. steamcoils) built into the regenerator itself, whereby one may extract heatfrom the catalyst during regeneration. Heat exchangers can be built intocatalyst transfer lines, such as for instance the catalyst return linefrom the regenerator to the reactor, whereby heat may be removed fromthe catalyst after it is regenerated. The amount of heat imparted to thecatalyst in the regenerator may be restricted by reducing the amount ofinsulation on the regenerator to permit some heat loss to thesurrounding atmosphere, especially if feeds of exceedingly high cokingpotential are planned for processing; in general, such loss of heat tothe atmosphere is considered economically less desirable than certain ofthe other alternatives set forth herein. One may also inject coolingfluids into the regenerator, for example water and/or steam, whereby theamount of inert gas available in the regenerator for heat absorption andremoval is increased.

Another suitable and preferred technique for controlling or restrictingthe heat transmitted to fresh feed via recycled regenerated catalystinvolves maintaining a specified ratio between the carbon dioxide andcarbon monoxide formed in the regenerator while such gases are in heatexchange contact or relationship with catalyst undergoing regeneration.In general, all or a major portion by weight of the coke present on thecatalyst immediately prior to regeneration is removed in at least onecombustion zone in which the aforesaid ratio is controlled as describedbelow. More particularly, at least the major portion more preferably atleast about 65% and more preferably at least about 80% by weight of thecoke on the catalyst is removed in a combustion zone in which the molarratio of CO₂ to CO is maintained at a level substantially below 5, e.g.about 4 or less. Looking at the CO₂ /CO relationship from the inversestandpoint, it is preferred that the CO/CO₂ molar ratio should be atleast about 0.25 and preferably at least about 0.3 and still morepreferably about 1 or more or even 1.5 or more. While persons skilled inthe art are aware of techniques for inhibiting the burning of CO to CO₂,it has been suggested that the mole ratio of CO:CO₂ should be kept lessthan 0.2 when regenerating catalyst with large heavy metal accumulationsresulting from the processing of carbo-metallic oils; in this connectionsee for example U.S. Pat. No. 4,162,213 to Zrinscak, Sr. et al. In thisinvention however, maximizing CO productionwhile regenerating catalystto about 0.1% carbon or less, and preferably about 0.05% carbon or less,is a particularly preferred embodiment of this invention. Moreover,according to a preferred method of carrying out the invention thesub-process of regeneration, as a whole, may be carried out to theabove-mentioned low levels of carbon on regenerated catalyst with adeficiency of oxygen; more specifically, the total oxygen supplied tothe one or more stages of regeneration can be and preferably is lessthan the stoichiometric amount which would be required to burn allhydrogen in the coke to H₂ O and to burn all carbon in the coke to CO₂.If the coke includes other combustibles, the aforementionedstoichiometric amount can be adjusted to include the amount of oxygenrequired to burn them.

Still another particularly preferred technique for controlling orrestricting the regeneration heat imparted to fresh feed via recycledcatalyst involves the diversion of a portion of the heat borne byrecycled catalyst to added materials introduced into the reactor, suchas the water, steam, naphtha, other hydrogen donors, flue gases, inertgases, and other gaseous or vaporizable materials which may beintroduced into the reactor.

The larger the amount of coke which must be burned from a given weightof catalyst, the greater the potential for exposing the catalyst toexcessive temperatures. Many otherwise desirable and useful crackingcatalysts are particularly susceptible to deactivation at hightemperatures, and among these are quite a few of the costly molecularsieve or zeolite types of catalyst. The crystal structures of zeolitesand the pore structures of the catalyst carriers generally are somewhatsusceptible to thermal and/or hydrothermal degradation. The use of suchcatalysts in catalytic conversion processes for carbo-metallic feedscreates a need for regeneration techniques which will not destroy thecatalyst by exposure to highly severe temperatures and steaming. Suchneed can be met by a multistage regeneration process which includesconveying spent catalyst into a first regeneration zone and introducingoxidizing gas thereto. The amount of oxidizing gas that enters saidfirst zone and the concentration of oxygen or oxygen bearing gas thereinare sufficient for only partially effecting the desired conversion ofcoke on the catalyst to carbon oxide gases. The partially regeneratedcatalyst is then removed from the first regeneration zone and isconveyed to a second regeneration zone. Oxidizing gas is introduced intothe second regeneration zone to provide a higher concentration or oxygenor oxygen-containing gas than in the first zone, to complete the removalof carbon to the desired level. The regenerated catalyst may then beremoved from the second zone and recycled to the reactor for contactwith fresh feed. An example of such multi-stage regeneration process isdescribed in U.S. Patent Application Ser. No. 969,602 of George D. Myerset al, filed Dec. 14, 1978, the entire disclosure of which is herebyincorporated herein by reference. Another example may be found in U.S.Pat. No. 2,938,739.

Multi-stage regeneration offers the possibility of combining oxygendeficient regeneration with the control of the CO:CO₂ molar ratio. Thus,about 50% or more, more preferably about 65% to about 95%, and morepreferably about 80% to about 95% by weight of the coke on the catalystimmediately prior to regeneration may be removed in one or more stagesof regeneration in which the molar ratio of CO:CO₂ is controlled in themanner described above. In combination with the foregoing, the last 5%or more, or 10% or more by weight of the coke originally present, up tothe entire amount of coke remaining after the preceding stage or stages,can be removed in a subsequent stage of regeneration in which moreoxygen is present. Such process is susceptible of operation in such amanner that the total flue gas recovered from the entire, completedregeneration operation contains little or no excess oxygen, i.e. on theorder of about 0.2 mole percent or less, or as low as about 0.1 molepercent or less, which is substantially less than the 2 mole percentwhich has been suggested elsewhere. Thus, multi-stage regeneration isparticularly beneficial in that it provides another convenient techniquefor restricting regeneration heat transmitted to fresh feed viaregenerated catalyst and/or reducing the potential for thermaldeactivation, while simultaneously affording an opportunity to reducethe carbon level on regenerated catalyst to those very low percentages(e.g. about 0.1% or less) which particularly enhance catalyst activity.Moreover, where the regeneration conditions, e.g. temperature oratmosphere, are substantially less severe in the second zone than in thefirst zone (e.g. by at least about 10 and preferably at least about 20°F.), that part of the regeneration sequence which involves the mostsevere conditions is performed while there is still an appreciableamount of coke on the catalyst. Such operation may provide someprotection of the catalyst from the more severe conditions. Aparticularly preferred embodiment of the invention is two-stagefluidized regeneration at a maximum temperature of about 1500° F. with areduced temperature of at least about 10° or 20° F. in the dense phaseof the second stage as compared to the dense phase of the first stage,and with reduction of carbon on catalyst to about 0.5% or less or evenabout 0.025% or less by weight in the second zone. In fact, catalyst canreadily be regenerated to carbon levels as low as 0.01% by thistechnique, even though the carbon on catalyst prior to regeneration isas much as about 1%.

In most circumstances, it will be important to insure that no adsorbedoxygen containing gases are carried into the riser by recycled catalyst.Thus, whenever such action is considered necessary, the catalystdischarged from the regenerator may be stripped with appropriatestripping gases to remove oxygen containing gases. Such stripping mayfor instance be conducted at relatively high temperatures, for exampleabout 1350° to about 1370° F., using steam, nitrogen or other inert gasas the stripping gas(es). The use of nitrogen and other inert gases isbeneficial from the standpoint of avoiding a tendency towardhydro-thermal catalyst deactivation which may result from the use ofsteam.

The following comments and discussion relating to metals management,carbon management and heat management may be of assistance in obtainingbest results when operating the invention. Since these remarks are forthe most part directed to what is considered the best mode of operation,it should be apparent that the invention is not limited to theparticular modes of operation discussed below. Moreover, since certainof these comments are necessarily based on theoretical considerations,there is no intention to be bound by any such theory, whether expressedherein or implicit in the operating suggestions set forth hereinafter.

Although discussed separately below, it is readily apparent that metalsmanagement, carbon management and heat management are inter-related andinterdependent subjects both in theory and practice. While coke yieldand coke laydown on catalyst are primarily the result of the relativelylarge quantities of coke precursors found in carbo-metallic oils, theproduction of coke is exacerbated by high metals accumulations, whichcan also significantly affect catalyst performance. Moreover, the degreeof success experienced in metals management and carbon management willhave a direct influence on the extent to which heat management isnecessary. Moreover, some of the steps taken in support of metalsmanagement have proved very helpful in respect to carbon and heatmanagement.

As noted previously the presence of a large heavy metals accumulation onthe catalyst tends to aggravate the problem of dehydrogenation andaromatic condensation, resulting in increased production of gases andcoke for a feedstock of a given Ramsbottom carbon value. Theintroduction of substantial quantities of H₂ O into the reactor, eitherin the form of steam or liquid water, appears highly beneficial from thestandpoint of keeping the heavy metals in a less harmful form, i.e. theoxide rather than metallic form. This is of assistance in maintainingthe desired selectivity.

Also, a unit design in which system components and residence times areselected to reduce the ratio of catalyst reactor residence time relativeto catalyst regenerator residence time will tend to reduce the ratio ofthe times during which the catalyst is respectively under reductionconditions and oxidation conditions. This too can assist in maintainingdesired levels of selectivity.

Whether the metals content of the catalyst is being managed successfullymay be observed by monitoring the total hydrogen plus methane producedin the reactor and/or the ratio of hydrogen to methane thus produced. Ingeneral, it is considered that the hydrogen to methane mole ratio shouldbe less than about 1 and preferably about 0.6 or less, with about 0.4 orless being considered about optimum.

Careful carbon management can improve both selectivity, (the ability tomaximize production of valuable products) and heat productivity. Ingeneral, the techniques of metals control described above are also ofassistance in carbon management. The usefulness of water addition inrespect to carbon management has already been spelled out inconsiderable detail in that part of the specification which relates toadded materials for introduction into the reaction zone. In general,those techniques which improve dispersion of the feed in the reactionzone should also prove helpful; these include for instance the use offogging or misting devices to assist in dispersing the feed.

Catalyst to oil ratio is also a factor in heat management. In commonwith prior FCC practice on VGO, the reactor temperature may becontrolled in the practice of the present invention by respectivelyincreasing or decreasing the flow of hot regenerated catalyst to thereactor in response to decreases and increases in reactor temperature,typically the outlet temperature in the case of a riser type reactor.Where the automatic controller for catalyst introduction is set tomaintain an excessive catalyst to oil ratio, one can expectunnecessarily large rates of carbon production and heat release,relative to the weight of fresh feed charged to the reaction zone.

Relatively high reactor temperatures are also beneficial from thestandpoint of carbon management. Such higher temperatures foster morecomplete vaporization of feed and disengagement of product fromcatalyst.

Carbon management can also be facilitated by suitable restriction of thetotal pressure in the reactor and the partial pressure of the feed. Ingeneral, at a given level of conversion, relatively small decreases inthe aforementioned pressures can substantially reduce coke production.This may be due to the fact that restricting total pressure tends toenhance vaporization of high boiling components of the feed, encouragecracking and facilitate disengagement of both unconverted feed andhigher boiling cracked products from the catalyst. It may be ofassistance in this regard to restrict the pressure drop of equipmentdownstream of and in communication with the reactor. But if it isdesired or necessary to operate the system at higher total pressure,such as for instance because of operating limitations (e.g. pressuredrop in downstream equipment) the above described benefits may beobtained by restricting the feed partial pressure. Suitable ranges fortotal reactor pressure and feed partial pressure have been set forthabove, and in general it is desirable to attempt to minimize thepressures within these ranges.

The abrupt separation of catalyst from product vapors and unconvertedfeed (if any) is also of great assistance. It is for this reason thatthe so-called vented riser apparatus and technique disclosed in U.S.Pat. Nos. 4,070,159 and 4,066,533 to George D. Myers et al is thepreferred type of apparatus for conducting this process. For similarreasons, it is beneficial to reduce insofar as possible the elapsed timebetween separation of catalyst from product vapors and the commencementof stripping. The vented riser and prompt stripping tend to reduce theopportunity for coking of unconverted feed and higher boiling crackedproducts absorbed on the catalyst.

A particularly desirable mode of operation from the standpoint of carbonmanagement is to operate the process in the vented riser using ahydrogen donor if necessary, while maintaining the feed partial pressureand total reactor pressure as low as possible, and incorporatingrelatively large amounts of water, steam and if desired, other diluents,which provide the numerous benefits discussed in greater detail above.Moreover, when liquid water, steam, hydrogen donors, hydrogen and othergaseous or vaporizable materials are fed to the reaction zone, thefeeding of these materials provides an opportunity for exercisingadditional control over catalyst to oil ratio. Thus, for example, thepractice of increasing or decreasing the catalyst to oil ratio for agiven amount of decrease or increase in reactor temperature may bereduced or eliminated by substituting either appropriate reduction orincrease in the charging ratios of the water, steam and other gaseous orvaporizable material, or an appropriate reduction or increase in theratio of water to steam and/or other gaseous materials introduced intothe reaction zone.

Heat management includes measures taken to control the amount of heatreleased in various parts of the process and/or for dealing successfullywith such heat as may be released. Unlike conventional FCC practiceusing VGO, wherein it is usually a problem to generate sufficient heatduring regeneration to heat balance the reactor, the processing ofcarbo-metallic oils generally produces so much heat as to requirecareful management thereof.

Heat management can be facilitated by various techniques associated withthe materials introduced into the reactor. Thus, heat absorption by feedcan be maximized by minimum preheating of feed, it being necessary onlythat the feed temperature be high enough so that it is sufficientlyfluid for successful pumping and dispersion in the reactor. When thecatalyst is maintained in a highly active state with the suppression ofcoking (metals control), so as to achieve higher conversion, theresultant higher conversion and greater selectivity can increase theheat absorption of the reaction. In general, higher reactor temperaturespromote catalyst conversion activity in the face of more refractory andhigher boiling constituents with high coking potentials. While the rateof catalyst deactivation may thus be increased, the higher temperatureof operation tends to offset this loss in activity. Higher temperaturesin the reactor also contribute to enhancement of octane number, thusoff-setting the octane depressant effect of high carbon lay down. Othertechniques for absorbing heat have also been discussed above inconnection with the introduction of water, steam, and other gaseous orvaporizable materials into the reactor.

The severe stripping and various regeneration techniques discussed aboveare useful in controlling heat release in the regenerator. While removalof heat from catalyst in or downstream of the regenerator by means ofheat exchangers (including steam coils) has been suggested as a meansfor controlling heat release, the above described techniques ofmulti-stage regeneration and control over the CO/CO₂ ratio (in eithersingle or multi-stage regeneration) are considered more advantageous.The use of steam coils is considered to be partly self-defeating, inthat a steam coil or heat exchanger in the regenerator or catalystreturn line will generally cause an increase in the catalyst to oilratio with a resultant increase in the rates of carbon production in thereactor and heat release in the regenerator.

ILLUSTRATIVE EMBODIMENTS

As noted above, the invention can be practiced in the above describedmodes and many others. Two illustrative, nonlimiting examples aredescribed by the accompanying schematic diagrams in FIGS. 1 and 2 and bythe description of those figures which follow.

FIG. 1 is a schematic diagram of an apparatus for carrying out theprocess of the present invention. The carbometallic oil feed (which mayhave been heated in a feed preheater not shown) and water suppliedthrough delivery pipe 9, are fed by feed supply pipe 10 having a controlvalve 11 to a wye 12 in which they mix with a flow of catalyst deliveredthrough supply pipe 13 and controlled by valve 14. Of course a varietyof mixing arrangements may be employed, and provisions may be made forintroducing the other added materials discussed above. The mixture ofcatalyst and feed, with or without other additional materials, is thenintroduced into riser 18.

Although riser 18 appears vertical in the drawing, persons skilled inthe art will recognize that the riser need not be vertical, as risertype reactors are known in which an appreciable portion of the riserpipe is non-vertical. Thus, riser pipes having an upward component ofdirection are contemplated, and usually the upward component of theirupwardly flowing inclined portions is substantial, i.e., at least about30°. It is also known to provide risers which have downwardly flowinginclined or vertical portions, as well as horizontal portions. Foldedrisers are also known, in which there are both upwardly extending anddownwardly extending segments. Moreover, it is entirely feasible topractice the process of the invention in an inclined and/or verticalpipe in which the feed and catalyst are introduced at an upper elevationand in which the feed and catalyst moves under the influence of gravityand the down flow of the feed to a lower elevation. Thus, in general,the invention contemplates the use of reaction chambers having a longL/D ratio and having a significant deviation from horizontal.

At the upper end of the riser 18 is a chamber 19 which receives thecatalyst from the riser. While chamber 19 may be a conventionaldisengagement and collection chamber, it is preferred that means beprovided for causing product vapors to undergo a sufficient change ofdirection relative to the direction traveled by the catalyst particles,whereby the vapors are suddenly and effectively separated from thecatalyst. Preferably, there is "ballistic" separation of catalystparticles and product vapors as described above.

In the present schematic diagram, the disengagement chamber 19 includesan upward extension 20 of riser pipe 18 having an open top 21 throughwhich the catalyst particles are discharged. This embodiment makes useof the so-called vented riser described in the above-mentioned Myers etal patents. Because of the refractory nature of carbo-metallicfractions, relatively high severity is required, but the rapiddisengagement of catalyst from lighter cracked products in the ventedriser prevents overcracking of desirable liquid products such asgasoline to gaseous products. The product vapors are caused to undergo asudden change of direction into lateral port 22 in the side of riserextension 20, the catalyst particles being, for the most part, unable tofollow the product vapors into port 22.

The vapors and those few particles which do manage to follow them intoport 22 are transferred by cross pipe 23 to a cyclone separator 24. Itis an advantage of the vented riser system shown that it can functionsatisfactorily with a single stage cyclone separator. However, in thepresent embodiment the cyclone separator 24 is employed as a first stagecyclone separator which is connected via transfer pipe 17 with optionalsecondary cyclone separator 25. The cyclone separator means, whether ofthe single- or multi-stage type, separates from the product vapors thosesmall amounts of catalyst which do enter the lateral port 22. Productvapors are discharged from disengagement chamber 19 through productdischarge pipe 26.

The catalyst particles which discharge from open top 21 of riser pipeextension 20, and those catalyst particles which are discharged from thedischarge legs 27 and 28 of primary and secondary cyclones 24 and 25drop to the bottom of disengagement chamber 19. The inventory andresidence time of catalyst in chamber 19 are preferably minimized.During startup those catalyst particles which are present may be kept insuspension by fluffing jets 30 supplied with steam through steam supplypipe 29. Spent catalyst spilling over from the bottom of disengagementchamber 19 passes via drop leg 31 to a stripper chamber 32 equipped withbaffles 33 and steam jet 34. Any of the other stripping gases referredto above may be employed with or in place of the steam.

Carbon is burned from the surface of the catalyst in the combustor 38which receives stripped catalyst via downcomer pipe 39 and control valve40. Blowers 41 and 42, in association with a valve and pipingarrangement generally indicated by 44, supply air to combustion air jets48 at the bottom of the combustor and to fluffing jets 49 at an elevatedposition. Air preheater 43, although usually unused when processingheavy hydrocarbons in accordance with the invention, may be employedwhen starting up the unit on VGO; then, when the unit is switched overto the carbo-metallic feed, preheater operation may be discontinued (orat least reduced). Supplemental fuel means may be provided to supplyfuel through the combustion air jets 48; but such is usually unnecessarysince the carbon lay down on the catalyst supplies more than enough fuelto maintain the requisite temperatures in the regeneration section.Regenerated catalyst, with most of the carbon burned off, departs thecombustor through an upper outlet 50 and cross pipe 51 to a secondarychamber 52, where it is deflected into the lower portion of the chamberby a baffle 53. Although the use of two stage regeneration iscontemplated, and preferred, in this particular embodiment the secondarychamber 52 was operated primarily as a separator chamber, although itcan be used to remove additional carbon down to about 0.01% or less inthe final stages of regeneration.

Catalyst moves in up to three different directions from the secondarychamber 52. A portion of the catalyst may be circulated back tocombustor 38 via catalyst recirculation loop 55 and control valve 56 forheat control in the combustor. Some of the catalyst is entrained in theproduct gases, such as CO and/or CO₂ produced by burning the carbon onthe catalyst in the combustor, and the entrained catalyst fines passupwardly in chamber 52 to two sets of primary and secondary cyclonesgenerally indicated by 57 and 58 which separate these catalyst finesfrom the combustion gases. Catalyst collected in the cyclones 57,58 anddischarged through their drop legs is directed to the bottom of chamber52 where catalyst is kept in suspension by inert gas and/or steam jets59 and by a baffle arrangement 54, the latter facilitating discharge ofregenerated catalyst through outlet 69 to catalyst supply pipe 13through which it is recirculated for contact with fresh feed at wyte 12,as previously described.

Combustion product gases produced by regeneration of the catalyst andseparated from entrained catalyst fines by the sets 57,58 of primary andsecondary cyclones in chamber 52, discharge through effluent pipes 61,62and heat exchangers 60,63. If such gases contain significant amounts ofCO, they may be sent via gas supply pipe 64 to an optional furnace 65 inwhich the CO is burned to heat heating coil 66 connected with steamboiler 67. Additional heat may be supplied to the contents of theboilers through conduit loop 68, which circulates fluid from the boiler67 to heat exchangers 60,63 and back to the boiler. This is of courseonly one example of many possible regeneration arrangements which may beemployed. The amount of heat passed from the regenerator back to theriser via regenerated catalyst may be controlled in any of the otherways described above; however it is preferred to control the relativeproportions of carbon monoxide and carbon dioxide produced while thecatalyst is in heat exchange relationship with the combustion gasesresulting from regeneration. An example of this technique is disclosedin the particularly preferred embodiment described in FIG. 2.

In FIG. 2 reference numeral 80 identifies a feed control valve infeedstock supply pipe 82. Supply pipe 83 introduces liquid water intothe feed. Heat exchanger 81 in supply pipe 82 acts as a feed preheater,whereby preheated feed material may be delivered to the bottom of risertype reactor 91. Catalyst is delivered to the reactor through catalyststandpipe 86, the flow of catalyst being regulated by a control valve 87and suitable automatic control equipment (not shown) with which personsskilled in the art of designing and operating riser type crackers arefamiliar.

The riser 91 may optionally include provision for injection of steamand, if desired, other gaseous and/or vaporizable material for thepurpose described above. The reactor is equipped with a disengagementchamber 92 similar to the disengagement chamber 19 of the reactor shownin FIG. 1, and the FIG. 2 embodiment thus includes means for causingproduct vapors to undergo a change of direction for sudden and effectiveseparation from the catalyst as in the previous embodiment. Catalystdeparts disengagement chamber 92 through stripper 94 which operates in amanner similar to stripper 32 of FIG. 1. Spent catalyst passes fromstripper 94 to regenerator 101 via spent catalyst transfer pipe 97having a slide valve 98 for controlling the flow.

Regenerator 101 is divided into upper chamber 102 and lower chamber 103by a divider panel 104 intermediate the upper and lower ends of theregenerator. The spent catalyst from transfer pipe 97 enters upperchamber 102 in which the catalyst is partially regenerated. Afunnel-like collector 106 having a bias-cut upper edge receivespartially regenerated catalyst from the upper surface of the dense phaseof catalyst in upper chamber 102 and delivers it via drop leg 107 havingan outlet 110 beneath the upper surface of the dense phase of catalystin lower regeneration chamber 103. Instead of the internal catalyst dropleg 107, one may use an external drop leg. Valve means in such externaldrop leg can control the catalyst residence time and flow rate in andbetween the upper and lower chambers.

Air is supplied to the regenerator through an air supply pipe 113. Aportion of the air travels through a branch supply pipe 114 to bayonet115 extending upwardly in the interior of plenum 111 along its centralaxis. Catalyst in chamber 103 has access to the space within plenum 11between its walls and the bayonet 115. A small bayonet (not shown) inthe aforementioned space fluffs the catalyst and urges it upwardlytoward a horizontally arranged ring distributor (not shown) where theopen top of plenum 111 opens into chamber 103. The remainder of the airpassing through air supply pipe 113 may be heated in air heater 117 (atleast during start-up with VGO) and is then introduced into the inlet118 of the aforementioned ring distributor, which may be provided withholes, nozzles or other apertures which produce an upward flow of gas tofluidize the partially regenerated catalyst in chamber 103.

The air introduced in the manner described above completes in chamber103 the regeneration of the partially regenerated catalyst received viadrop leg 107. The amount of air that is supplied is sufficient so thatthe air and/or the resultant combustion gases are still able to supportcombustion upon reaching the top of chamber 103. The aforementioned dropleg 107 extends through an enlarged aperture in panel 104, to which issecured a gas distributor 120 which is concentric with and surrounds thedrop leg. Via gas distributor 120, combustion supporting gases, whichhave now been partially depleted of combustion supporting gas, areintroduced into the upper regenerator chamber 102 where they contact forpurposes of partial oxidation the incoming spent catalyst from spentcatalyst transfer pipe 97. Apertured probes 121 or other suitable meansin gas distributor 120 assist in achieving a uniform distribution of thepartially depleted combustion supporting gas in upper chamber 102.Supplemental air or other fluids may be introduced into upper chamber102, if desired through supply pipe 122, which discharges into orthrough gas distributor 120.

Fully regenerated catalyst with less than about 0.25% carbon, preferablyless than about 0.1% and more preferably less than about 0.05%, isdischarged from lower regenerator chamber 103 through a regeneratedcatalyst stripper 128, whose outlet feeds into the catalyst standpipe 86mentioned above. Thus, regenerated catalyst is returned to riser 91 forcontact with additional fresh feed from feed supply pipe 82. Whateverheat is introduced into the recycled catalyst in the regenerator 101 isavailable for heat transfer with the fresh feed in the riser.

The division of the regenerator into upper and lower regenerationchambers 102 and 103 not only smooths out variations in catalystregenerator residence time but is also uniquely of assistance inrestricting the quantity of regeneration heat which is imparted to thefresh feed while yielding a regenerated catalyst with low levels ofresidual carbon for return to the reactor. Because of the arrangement ofthe regenerator, the spent catalyst from transfer line 97, with its highloading of carbon, contacts in chamber 102 combustion supporting gaseswhich have already been at least partially depleted of oxygen by theburning of carbon from partially regenerated catalyst in lowerregenerator chamber 102. Because of this, it is possible to control boththe combustion and the quantity of carbon dioxide produced in upperregenerator chamber 102. Although the air or other regenerating gasintroduced through air supply pipe 113 and branch conduit 114 maycontain a relatively large quantity of oxygen, the partially regeneratedcatalyst which they contact in lower regenerator chamber 103 has alreadyhad part of its carbon removed. The high concentration of oxygen and thetemperature of the partially regenerated catalyst combine to rapidlyremove the remaining carbon in the catalyst, thereby achieving a cleanregenerated catalyst with a minimum of heat release. Thus, here again,the combustion temperature and the CO:CO₂ ratio in the lowerregeneration chamber are therefore readily controlled. The regenerationoff gases are discharged from upper regenerator chamber 102 via off gaspipe 123, regulator valve 124, catalyst fines trap 125 and outlet 126.

The vapor products from disengagement chamber 92 may be processed in anyconvenient manner such as for example, by discharge through vapordischarge line 131 to the inlet of fractionator 132. Said fractionatorincludes a bottoms outlet 133, side outlet 134, flush oil stripper 135,and stripper bottom outlet and discharge line 136 connected to pump 137for discharging flush oil. The overhead product from stripper 135 isrouted via stripper overhead return line 138 to the fractionator 132.

The main overhead discharge line 139 of the fractionator is connected tooverhead receiver 142 having a bottoms discharge line 143 feeding intopump 144 for discharging gasoline product. If desired, a portion of thisproduct may be sent via recirculation line 145, the flow beingcontrolled by recirculation valve 146, back to the fractionator 132. Theoverhead receiver also includes a water receiver 147 and a waterdischarge line 148. The gas outlet 150 of the overhead receiverdischarges a stream which is mainly below C₅, but containing some C₅, C₆and C₇ material. If desired, the C₅ and above material in this gasstream may be separated by compression, cooling and fractionation andrecycled to the overhead receiver with a compressor, cooler andfractionator (not shown).

EXAMPLES

The following examples are given only by way of illustration and not forlimiting the invention. The properties of the feedstocks and catalystsemployed in the examples are set forth in tables I and II above.

EXAMPLES 1-3

A unit constructed in accordance with FIG. 1 was operated in the mannerand with the results set forth in Table III and in accordance with theteachings of FIG. 1 and the associated description thereof.

EXAMPLES 4

This example was conducted as indicated in Table III with a unitconstructed in accordance with the teachings of FIG. 2. Regeneration wascarried out in two stages.

                  TABLE III                                                       ______________________________________                                                  EXAMPLES                                                                                               4                                                    1      2        3        ABL R.C..sup.2                             ______________________________________                                        Feedstock                                                                     Feed #/hr..sup.1                                                                          75,215   74,478   72,209 2268                                     Steam #/hr. 4,050    3,840    3,910  193.0                                    Water #/hr. 8,210    9,230    9,380  96.4                                     Catalyst    GRZ-1    GRZ-1    GRZ-1                                           Cat/Oil Ratio                                                                             6.0      5.14     5.8    9.6                                      Riser Vapor                                                                   Residence Time                                                                (Seconds)                                                                     Rx Temperature -                                                                          1010     1012     1010   980                                      °F.                                                                    Pressure (psia)                                                                           35       36.3     34.6   35.1                                     Products                                                                      Dry Gas Wt. %                                                                             6.1      4.9      5.6    4.8                                      C.sub.3 Sat. Vol. %                                                                       3.5      3.2      2.2    2.7                                      C.sub.3 Olefins Vol. %                                                                    8.9      5.4      5.9    9.2                                      C.sub.4 Sat. Vol. %                                                                       3.9      6.8      3.5    7.0                                      C.sub.4 Olefins Vol. %                                                                    7.1      7.3      8.3    9.7                                      C.sub.5 - 430° Vol. %                                                              46.1     33.4     45.2   51.7                                     Coke Wt. %  10.0     8.9      10.0   11.6                                     Material Balance                                                                          105.8    97.4     102.7  95.1                                     Wt. %                                                                         Selectivity .72      .58      .70    .63                                      Conversion, Vol. %                                                                        63.6     59.8     64.3   82.1                                     Volume Yield                                                                              123.6    111.3    118.1  111.0                                    Liquid Vol. %                                                                 Gasoline                                                                      API Gravity 54.2     53.0     52.7   47.8                                     Distillation                                                                  IBP         94       94       104    168                                      50%         238      254      255    264                                      95%         434      438      435    420                                      RONC        91.6     89.4     89.2   92.3                                     MONC        80.6     78.7     78.9   82.7                                     Regeneration                                                                              1350     1336     1330   N/A                                      Temp. °F.                                                              Upper Regen.                                                                              N/A      N/A      N/A    1303                                     Temp. °F                                                               Lower Regen.                                                                              N/A      N/A 1    N/A    1325                                     Temp. °F.                                                              Spent Catalyst                                                                            1.67     1.72     1.72   1.17                                     Wt. % C                                                                       Regenerated .43      .53      .53    .01                                      Catalyst % C                                                                  Circulating                          .34                                      Catalyst % C                                                                  O.sub.2 Effluent                                                                          0        0        0      0                                        Mol % of                                                                      total flue gas                                                                CO.sub.2 /CO lb./lb.                                                                      2.67     2.43     3.19   3.81                                     ______________________________________                                         .sup.1 Fresh feed only, except that Example 3 includes recycle of 9,477       #/hr. of HCO (heavy cycle oil), having an API gravity of 18.3 and boiling     at about 550 to 775, injected into the reduced crude line to the riser.       .sup.2 Arabian Light Reduced Crude                                       

What is claimed is:
 1. A process for economically convertingcarbometallic oils to lighter products, comprising:I. providing aconverter feed containing 650° F.+ material, said 650° F.+ materialbeing characterized by a carbon residue on pyrolysis of at least about 1and by containing at least about 4 parts per million of NickelEquivalents of heavy metal(s); II. bringing said converter feed togetherwith cracking catalyst having an equilibrium microactivity testconversion activity level of at least about 40 and bearing anaccumulation of at least about 3000 ppm Nickel Equivalents of heavymetal(s) expressed as metal(s) by weight on regenerated equilibriumcatalyst; III. bringing said converter feed together with liquid waterin a weight ratio relative to feed in the range of about 0.04 to about0.15; IV. forming a stream containing a mixture of said converter feed,said catalyst and steam resulting from the vaporization of said liquidwater and causing the resultant stream to flow through a progressiveflow type reactor having an elongated reaction chamber which is at leastin part vertical or inclined at a lineal velocity of at least about 25feet per second for a vapor residence time in the range of about 0.5 toabout 10 seconds at a reaction chamber outlet temperature of about 975°to about 1300° F. and under a pressure of about 10 to about 50 poundsper square inch absolute sufficient for causing a conversion per pass inthe range of about 60% to about 90% while producing coke in amounts inthe range of about 6 to about 14% by weight based on fresh feed, andlaying down coke on the catalyst in amounts in the range of about 0.3 toabout 3% by weight; V. separating said catalyst from the catalystcracking products; VI. stripping said separated catalyst; VII.regenerating said catalyst; and VIII. recycling the regenerated catalystto the reactor for contact with fresh feed.
 2. A process according toclaim 1 in which said 650° F.+ material represents at least about 70% byvolume of said feed.
 3. A process according to claim 1 in which the 650°F.+ material includes at least about 10% by volume of material whichwill not boil below about 1000° F.
 4. A process according to claim 1 inwhich the 650° F.+ material includes at least about 10% by volume ofmaterial which will not boil below about 1025° F.
 5. A process accordingto claim 1 wherein said carbon residue corresponds with a Ramsbottomcarbon value in the range of about 2 to about
 12. 6. A process accordingto claim 1 wherein the carbon residue of the feed as a whole correspondswith a Ramsbottom carbon value of at least about
 1. 7. A processaccording to claim 6 wherein said Ramsbottom carbon value is in therange of about 2 to about
 12. 8. A process according to claim 1 whereinthe feed as a whole contains at least about 4 parts per million ofNickel Equivalents of heavy metal(s), of which at least about 2 partsper million is nickel (expressed as metal(s), by weight).
 9. A processaccording to claim 1 wherein at least about 85% by volume of the feed isoil which has not previously been contacted with cracking catalyst undercracking conditions.
 10. A process according to claim 1 wherein saidfeed comprises about 15% or less by volume of recycled oil.
 11. Aprocess according to claim 1 wherein said feed is processed in asubstantially once-through or single pass mode with no substantialamount of recycled oil in the feed.
 12. A process according to claim 1wherein said catalyst is contacted with said feed in said elongatedreaction zone in a weight ratio of catalyst to converter feed in therange of about 3 to about
 18. 13. A process according to claim 1 inwhich the liquid water is brought together with the converter feed at orprior to bringing converter feed into contact with the crackingcatalyst.
 14. A process according to claim 1 in which liquid water andconverter feed are mixed in an atomizing nozzle and sprayed into contactwith the catalyst.
 15. A process according to claim 1 wherein catalystis added to the process at a rate in the range of about 0.1 to about 3pounds per barrel of feed.
 16. A process according to claim 1 whereinsaid catalyst as introduced into the process has a microactivity of atleast about 60 volume percent.
 17. A process according to claim 1wherein said catalyst has an equilibrium microactivity of at least about60 volume percent.
 18. A process according to claim 1 wherein saidcatalyst is equilibrium cracking catalyst which has previously been usedin a fluid catalytic cracking unit in which said catalyst was used forthe cracking of feed characterized by a carbon residue on pyrolysis ofless than 1 and by containing less than about 4 parts per million ofNickel Equivalents of heavy metal(s).
 19. A process according to claim 1wherein there is an accumulation of heavy metal(s) on said catalyst inthe range of about 3000 ppm to about 70,000 ppm of Nickel Equivalents,by weight, expressed as metal(s) on regenerated equilibrium catalyst.20. A process according to claim 1 wherein there is an accumulation ofheavy metal(s) on said catalyst in the range of about 4000 ppm to about50,000 ppm of Nickel Equivalents, by weight, expressed as metal(s) onregenerated equilibrium catalyst.
 21. A process according to claim 1wherein there is an accumulation of heavy metal(s) on said catalyst inthe range of about 5000 ppm to about 30,000 ppm of Nickel Equivalents,by weight, expressed as metal(s) on regenerated equilibrium catalyst.22. A process according to claim 1 wherein said catalyst ischaracterized by a pore structure for absorbing hydrocarbon moleculesand by reactive sites within or adjacent the pores.
 23. A processaccording to claim 12 wherein said catalyst is a zeolite containingcatalyst.
 24. A process according to claim 23 wherein said zeolitecontaining catalyst is a molecular sieve catalyst which includes atleast about 5% by weight of sieve.
 25. A process according to claim 1wherein the total amount of gaseous and/or vaporized material other thanconverter feed and resultant products which is present in said reactionzone is in a weight ratio, relative to converter feed, of up to about0.4.
 26. A process according to claim 1 wherein the total amount ofgaseous and/or vaporized material other than converter feed andresultant products which is present in said reaction zone is in a weightratio, relative to converter feed, in the range of about 0.02 to about0.4.
 27. A process according to claim 1 wherein the total amount ofgaseous and/or vaporized material other than converter feed andresultant products which is present in said reaction zone is in a weightratio, relative to converter feed, in the range of about 0.03 to about0.3.
 28. A process according to claim 1 wherein the total amount ofgaseous and/or vaporized material other than converter feed andresultant products which is present in said reaction zone is in a weightratio, relative to converter feed in the range of about 0.05 to about0.25.
 29. A process according to claim 1 wherein said reactor is a risertype reactor.
 30. A process according to claim 1, wherein said reactoris a vented riser type reactor.
 31. A process according to claim 1wherein said residence time of the converter feed and product vapors isin the range of about 0.5 to about 6 seconds.
 32. A process according toclaim 1 wherein said residence time of the converter feed and productvapors is in the range of about 1 to about 4 seconds.
 33. A processaccording to claim 1 wherein said residence time of the converter feedand product vapors is in the range of about 1.5 to about 3 seconds. 34.A process according to claim 1 wherein the ratio of the average catalystresidence time to vapor residence time is in the range of about 1 toabout
 5. 35. A process according to claim 1 wherein the ratio of averagecatalyst residence time to vapor residence time is in the range of about1 to about
 4. 36. A process according to claim 1 wherein the ratio ofaverage catalyst residence time to vapor residence time is in the rangeof about 1.2 to about
 3. 37. A process according to claim 1 wherein theratio of average catalyst residence time to vapor residence time is inthe range of about 1.2 to about
 2. 38. A process according to claim 1wherein the reactor pressure is in the range of about 15 to about 35psia.
 39. A process according to claim 1 wherein the feed partial ortotal pressure is in the range of about 3 to about 30 psia.
 40. Aprocess according to claim 1 wherein the feed partial or total pressureis in the range of about 7 to about 25 psia.
 41. A process according toclaim 1 wherein the feed partial or total pressure is in the range ofabout 10 to about 17 psia.
 42. A process according to claim 1 in whichthe weight ratio of liquid water to converter feed is about 0.04 toabout 0.1.
 43. A process according to claim 1 in which the weight ratioof liquid water to converter feed is about 0.05 to about 0.1.
 44. Aprocess according to claim 1 wherein the coke laydown is in the range ofabout 0.5 to about 3%.
 45. A process for economically convertingcarbometallic oils to lighter products, comprising:I. providing aconverter feed containing 650° F.+ material, said converter feed beingcharacterized by a carbon residue on pyrolysis of at least about 1 andby containing at least about 4 parts per million of Nickel Equivalentsof heavy metal(s); II. bringing said converter feed together with liquidwater in a weight ratio relative to converter feed in the range of about0.04 to about 0.15; III. bringing said converter feed together withcracking catalyst having an equilibrium MAT conversion activity level ofat least about 60 volume percent and bearing an accumulation of at leastabout 3000 ppm by weight of nickel equivalents of heavy metal(s)expressed as metal(s) on regenerated equilibrium catalyst; IV. forming astream containing a mixture of said converter feed, said catalyst andsteam resulting from the vaporization of said liquid water and causingthe resultant stream to flow through a progressive flow type reactorhaving an elongated reaction chamber which is at least in part verticalor inclined at a lineal velocity of at least about 35 feet per secondfor a vapor residence time in the range of about 0.5 to about 6 secondsat a reaction chamber outlet temperature of about 975 to about 1300° F.and under a pressure of about 10 to about 50 pounds per square inchabsolute sufficient for causing a conversion per pass in the range ofabout 60% to about 90% while producing coke in amounts in the range ofabout 6 to about 14% by weight based on fresh feed, and laying down cokeon the catalyst in amounts in the range of about 0.3 to about 3% byweight; V. separating said catalyst from at least a substantial portionof the stream comprising said catalyst, steam and resultant crackingproducts formed in the elongated reaction chamber; VI. stripping saidseparated catalyst; VII. regenerating said catalyst; and VIII. recyclingthe regenerated catalyst to the reactor for contact with fresh feed. 46.A process according to claim 45 wherein said weight ratio of liquidwater to converter feed is in the range of about 0.04 to about 0.1. 47.A process according to claim 45 wherein H₂ O is brought into contactwith said converter feed in said stream and/or prior to formation ofsaid stream in the form of said liquid water and in the form of steam ina weight ratio relative to feed in the range of about 0.01 to about0.25, the total H₂ O thus supplied not exceeding a weight ratio of about0.3 relative to converter feed.
 48. A process according to claim 45wherein said accumulation of heavy metal(s) on said catalyst is at leastabout 4000 ppm of Nickel Equivalents, by weight, expressed as metal(s)on regenerated equilibrium catalyst.
 49. A process according to claim 45wherein said accumulation of heavy metal(s) on said catalyst is at leastabout 5000 ppm of Nickel Equivalents, by weight, expressed as metal(s)on regenerated equilibrium catalyst.
 50. A process according to claim 45wherein said elongated reaction chamber outlet termperature is in therange of about 985° to about 1200° F.
 51. A process according to claim45 wherein said conversion is in the range of about 70% to about 85%.52. A process for economically converting carbo-metallic oils to lighterproducts, comprising:I. providing a converter feed that has hadsubstantially no prior hydrotreatment and contains at least about 70% byvolume of 650° F.+ material and at least about 10% by volume of materialwhich will not boil below about 1025° F., said converter feed beingcharacterized by a carbon residue on pyrolysis of at least about 2 andby containing at least about 5.5 parts per million of Nickel Equivalentsof heavy metal(s); II. bringing said converter feed together with liquidwater in a weight ratio relative to converter feed in the range of about0.04 to about 0.15; III. at a catalyst to fresh converter feed weightratio of at least about 6, bringing said converter feed together withhot cracking catalyst having an equilibrium MAT conversion activitylevel of at least about 60 volume percent and bearing an accumulation ofmore than about 5000 ppm by weight of Nickel Equivalents of heavymetal(s) expressed as metal(s) on regenerated equilibrium catalyst; IV.forming a stream containing fine droplets of said converter feed, saidcatalyst and steam resulting from the vaporization of said liquid waterand causing the resultant stream to flow through a progressive flow typereactor having an elongated riser reaction chamber which is at least inpart vertical or inclined at a lineal velocity of at least about 35 feetper second for a vapor residence time in the range of about 0.5 to about3 seconds at a reaction chamber outlet temperature of about 985° toabout 1200° F. and under a pressure of about 15 to about 35 pounds persquare inch absolute sufficient for causing a conversion per pass in therange of about 60% to about 90% while producing coke in amounts in therange of about 6 to about 14% by weight based on fresh feed, and layingdown coke on the catalyst in amounts in the range of about 0.3 to about3% by weight, but insufficient for fully vaporizing the converter feed;V. separating said catalyst from at least a substantial portion of thestream comprising said catalyst, steam and resultant cracking productsformed in the elongated reaction chamber; VI. stripping said separatedcatalyst; VII. regenerating said catalyst; and VIII. recycling theregenerated catalyst to the reactor for contact with fresh feed.